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COMPENDIUM
21 st Refinery Technology Meet
20 22 April, 2017
Convention Centre, Hotel Novotel
Visakhapatnam, India
Refining to Petrochemicals – The Way Ahead
Organised by
Centre for High Technology
&
Hindustan Petroleum Corporation Limited
INDEX
ORAL PRESENTATIONS
1. Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of
the Barrel
Nicoletta Panariti, Eni Refining & Marketing
Raju Chopra, Haldor Topsoe India Pvt. Ltd.
2. Bottom Of Barrel Upgrading-Future Refinery Configuration
Jayesh Patel, Indian Oil Corporation Limited, Mathura Refinery
3. Multi- zone Catalytic Cracking (MCC) - A new platform for Direct Cracking of Crude &
Opportunity Feedstock to Light Olefins & BTX
Sukumar Mandal, Manoj Yadav, Nilam Limbasiya, Vinod Rayan, Sampath Nerivetla,
Gopal. Ravichandran, Praveen Chinthala, Asit Kumar Das
Refining R&D Centre, MAB, Reliance Industries Limited, Jamnagar, Gujarat
4. Progress and Future of HDPE Technology
Shin-ichi Kojoh, Mitsui Chemicals, Inc., Japan
5. 2nd generation HyBRIMTM atalyst: I o atio to fulfill the refi er s ish
Michael Tinning T. Schmidt, Raju Chopra and Hussain Lokhandwala
Haldor Topsoe A/S
6. Operating experience with the New Continuous Catalyst Withdrawal System for
Improved RFCC Operations
Paresh Amin & Kate Hovey, Johnson Matthey Process Technologies
7. BoroCat™ – An Innovative Solution for Resid Fluid Catalytic Cracking (FCC) Units
Bilge Yilmaz, BASF Corporation, U.S.A
Lynne Tan Xin Lin, BASF South East Asia Pte Ltd, Singapore
Aaron Liew, BASF South East Asia Pte Ltd, Singapore
8. GT-BTX PluS®: Generating Aromatics from FCC Gasoline
Sachin Joshi, GTC Technology US, LLC, Houston, Texas
9. Low Temperature Catalytic Gasification of Petcoke in to Syngas
Sateesh Daggupati, Sachchit Majhi, Sukumar K. Mandal, Asit K. Das
Refining R&D, Reliance Industries Limited, Jamnagar, Gujarat, India
10. Aromatics Production from Stranded Streams in a Refinery: A Techno-Commercial
Perspective
Hillol Das, Sanjay Rajora, Manasi Patel, Essar Oil Ltd., Mumbai
11. Numaligarh Refinery embarks into Bio-Refinery from Bamboo feedstock -A Game
Changer
Geetali Kalita & Rupam Kumar Sarmah, Numaligarh Refinery Ltd, Assam
12. Petroleum/Petrochemical Integration at BPCL Kochi Refinery
George Paul, S. Ramanathan, Bharat Petroleum Corp. Ltd., Kochi Refinery
13. A New Innovation in Steam Methane Reforming - High Efficiency, Zero Steam Export,
Lower Fuel Consumption
Dr. Siddhartha Mukherjee, Air Liquide Global E&C Solutions India Private Limited
14. An Improved Design of Threaded Closures for, Screw Plug Heat Exchangers
Haresh K. Sippy, TEMA India Limited
15. HTRI SmartPMTM U lo ks Resear h Fi di gs, Cuts Refi eries A
Throughput Costs
Simon J. Pugh, Heat Transfer Research, Inc., Navasota, USA
ual E ergy a d
16. Aerobic Granular Biomass Technology - NEREDA® – A Sustainable Wastewater
Treatment Option
K. Yagna Prasad, VA Tech Wabag Ltd.
17. Refi ery s ater sustai a le assess e t: The need of hour
Jayant Kumar Joshi, Engineers India Limited
S.N. Sukhwal, Centre for High Technology
18. Process & Reliability improvement in Reverse Osmosis (RO) system by Innovative
Chemical Dosing Modification
Mahendra Singh Bhadauria, Rajesh Nandanwar and Anand Pratap Raghav
Bharat Oman Refineries Ltd, Bina
19. Energy efficiency improvement through innovation
Raju Chopra, Maninder Jit Singh, Sagar Shukla, Sachin Panwar, Haldor Topsoe India
20. Maximization of BS-VI MS at Paradip Refinery
Geethashree, Rahul Srivastava & Amal K. Roy,
Indian Oil Corporation Limited, Paradip Refinery
21. H2 re o ery fro MP Separators apor strea s of DHDT u it
Soumitro Majumdar, Malay Bagchi, K K Mandal
Indian Oil Corporation Limited, Panipat Refinery
22. Squeeze more value out of your FCC gasoline
Anuj Seth and Claus Brostrom Nielsen, Haldor Topsoe
23. Model-Based Utilities Optimization and Management
Terumi Okano, Sunil Patil, Jack Zhang, AspenTech
24. Application of High Pressure Breech-Lock (Screw Plug) Closure Heat Exchangers for
Refinery Services
Ugrasen Yadav and Amrendra Bet, TECHNIP India Limited
25. Prediction of Reformate Octane Barrels for change in Crude Mix using Simulation
Model at HPCL-Mumbai Refinery (MR)
Rohitashva Tewari, Ashok Golekar, Ashok Kumar, Hindustan Petroleum Corp. Ltd.,
Mumbai Refinery
26. Optimising Flooded Columns -Practical Experience
Mannu Jha, Nagarameshkumar Parimi, Anand A. Haradi, Shyam P. Kamath,
Mangalore Refinery and Petrochemicals Ltd.
27. Approach towards Abatement of VOC Emission at Paradip Refinery
J.R. Behera, Shashi Vardhan (DGM-TS), Indian Oil Corporation Limited, Paradip Refinery
28. Strategy in setting up a Petrochemical Industry in NE India and its journey towards
Global Competitiveness: A study
Pranjal ChangmaiPradeep Rawat, Pranjal Kumar Phukan, Subodh Kumar, Tamagna
Ghosh, Manish Kumar Binjola
Brahmaputra Cracker & Polymer Limited (BCPL)
29. Energy Efficiency and Reliability Improvements in Design of Grass-root CDU/VDU at
HPCL Visakh Refinery
Deepak Kumar Jha, Ashraf Jamal
Hindustan Petroleum Corporation Limited, HQO, Mumbai
30. Increasing Hydrogen Recovery in PSA through Adsorbent Replacement: A First-of-itskind Effort by HPCL
Arun Kuniyil, Prashant Mishra, T A Rajiv Kumar, S N Sheshachala, Dr. Peddy V C Rao, Dr.
N V Choudary, G Sriganesh
HPCL Green R&D Centre, Bengaluru
S Tazeem Abbas, Ashok P Golekar, S V Choulkar, Mandar A Joshi
HPCL, Mumbai Refinery
31. A statistical approach to tap hidden fuel saving potential in GTGs
Subhash Nandanwar, M.V.Borkar and Sardar Shaik
Hindustan Petroleum Corporation Limited - Visakh Refinery
32. Introduction of Hot Separator in Naphtha Hydrotreater
Anirban Ray, Debasis Sarma, Reliance Refinery Division
POSTER PRESENTATIONS
1. Innovative Way to Replace Internal Gasket of CCR Platforming Reactor Reduces Unit
Downtime and Set New Benchmark
Akhilesh Kumar saxena, Ashish kumar, Manish Pandey, Rakesh Sharma, Vikant Maithil
Deepesh Kumar
Bharat Oman Refineries Limited (BORL), Bina, Madhya Pradesh
2. Assessment of hydrogen unit revamp by heat exchange reformer
Prashant Parihar, Ravi Kumar Voolapalli,
Corporate Research and development Centre, Bharat Petroleum Corporation Limited
Pankaj Muley, Nitin Jawale, A. C. Prabhune, V. Suresh
Bharat Petroleum Corporation Limited, Mumbai Refinery
3. Implementation of Reliability Centered Maintenance (RCM) and Failure Mode and
Effects Analysis (FMEA) In Essar Oil
Jayanti Vagdoda, Aditya Trivedi, Essar Oil
4. Upgradation of Pitch from Slurry phase Resid Hydrocracking
Bhavesh Sharma, Kanuparthy Naga Raja, Peddy V C Rao and G. Sriganesh
HP Green R&D Centre, Hindustan Petroleum Corporation Limited
5. Deployment of Parallel Solve Automation in LP models for speedier Spot Crude oil
evaluation
Mridusmita Goswami, R. Jerold, Kabidas Mandal, S.N. Pandey
Corporate Optimisation, Indian Oil Corporation Ltd.
6. Utilization of low cost Coal/Petcoke as fuel for power generation in place of fuel oil
Mamoni Basumatary Mrinal Basumatary, Deepa Thapa, K. Mashruwala,
Indian Oil Corporation Ltd., Guwahati Refinery
7. Provision of Plate type of Heat Exchanger in ARU
P.E. Kishore Babu, M. Sankar, Indian Oil Corporation Ltd., Mathura Refinery
8. Opportunities & Challenges for Future Purified Terephthalic Acid (PTA) Plant and
its Integration with Refinery
Kamleshwar Luckwal, Indian Oil Corporation Ltd., Panipat Refinery
9. Optimizing operating parameters by predicting feed quality
Mousom Some, Puranjay Choudhury, Indian Oil Corporation Ltd., Panipat Refinery
10. Future Indian Refinery Configuration – SWOT analysis and metamorphic change
towards Energy & Petrochemicals from standard fuel complex
Srinivas Moturi, Nagaphani Kumar Ravuri, Mangalore Refinery and Petrochemicals Ltd.
11. Fluidized bed gasification: Technology status, Challenges and Modeling approaches
Ankit A. Jain and Ajay Gupta, Refining R&D, Reliance Industries Limited, Jamnagar
12. First Principle Approach to Debottleneck Propylene Recovery Unit to Produce Polymer
Grade Propylene
Mukesh Kumar Sharma, S.G.Venkatesh, RHQ (T), Indian Oil Corporation Ltd.
13. DeltaV Embedded Advanced Control
Emerson
14. On Site Technical Performance Analysis of a MicroTurbine
Parivesh Chugh, T. Nandakumar, GAIL (India) Ltd., NOIDA
15. Holistic Refinery well-being through efficient and robust Amine system
Ashok Kumar,Ashok P. Golekar, Satish D. Khedekar, Gaurav Vyas
Hindustan Petroleum Corporation Limited, Mumbai Refinery
16. Effect of ZSM-5 zeolite crystal size in Fluid Catalytic Cracking Reactions
Praveen Chinthala, Gopal Ravichandran and Asit K. Das
Refining R&D, Reliance Technology Group, RIL, Jamnagar, Gujarat
17. Recovery of hydrogen from the refinery off-gas streams: Application of membrane
technology
Nitin Somkuwar, Renny Andrew, Sonal Maheshwari, Gokak D.T.
Corporate R&D Centre, Bharat Petroleum Corporation Ltd.
18. Opportunity Crudes (High Calcium) – Processing challenges
Nagashyam Appalla, Reliance refinery division, RIL
19. Crude to Petrochemicals – A New Paradigm
Sanjay Rajora, Manasi Patel, Hillol Das, Essar Oil Ltd., Mumbai
20. HP Gelators – Organic Soft Materials – A Potential Solution for Oil Spillage
Remediation and Recovery
Chinthalapati Siva Kesava Raju, Bhaskar Pramanik, Kandanelli Ramesh, Mangala
Ramkumar, Raman Ravishankar, Peddy Venkat Chalapathi Rao & G. Sriganesh
Hindustan Petroleum Green R&D Center, Bangalore
21. Back Casting – A novel Approach using Integrated Planning
Ramandeep Singh, K. Mandal, S.N Pandey
Corporate Optimisation, Indian Oil Corporation Ltd.
22. Replacement of PTA Slurry Incinerator by Environmental Friendly Flaker at Panipat
Refinery
Indian Oil Corporation Ltd., Panipat Refinery
23. Corrosion Mitigation in Hydrotreaters by Employing Best Design & Operating Practices
Rakesh Gagat, Mukesh Kumar RHQ (T), Indian Oil Corporation Limited, New Delhi
24. Managing Redoil Fouling in Depropaniser Column
Mannu Jha, Nagarameshkumar Parimi, Anand A. Haradi, Shyam P. Kamath
Mangalore Refinery and Petrochemicals Ltd.
25. Steam Coil Bypass in Out Board Steam Generator In CCRU
G.L. Arun Kumar, Appalaraju Pentakota, V. Venkatesh, S. Sankaran
Chennai Petroleum Corporation Limited
26. Microcrystalline waxes from Industrial polyolefinic by product
Manisha Sahai, Ajay Kumar and Sanat Kumar
CSIR-Indian Institute of Petroleum
27. Lead –Lag arrangement in Cycle Water Treating Unit (DI-Unit) in MEG Plant
Utsav Shankar, Syed Ashfaque Ali, Indian Oil Corporation Ltd., Panipat Refinery
28. Raising the bar of TAN Limitation of LK Fraction for ATF Production
Biswajit Shown, Swapan Ghosh and Asit Kumar Das
Reliance Industries Ltd., Jamnagar, Gujarat, India
29. A process for the production of Low Poly-Aromatic Hydrocarbon (PAH) Rubber Process
Oil (RPO) from Distillate Aromatic Extract (DAE) and other blending streams – Value
upgradation
V. Selvavathi, M. Lavanya and R. Krishnamurthy, Chennai Petroleum Corp. Ltd.
30. Improvements carried out in Oil Movement and Storage section to augment capacity
of HPCL Mumbai Refinery
A.B.Chattopadhyay, Ashok Kumar, Ganeshraj G, Syed Arif Hussain
Hindustan Petroleum Corporation Limited, Mumbai Refinery
31. Analysis of complex flows in a petroleum refinery
Muffazal Badshahwala, Rahul C Patil, Ajay Gupta, Asit Das
Reliance Industries Limited
32. Motive Steam pressure optimization via Vacuum Ejector Modelling
Mayur Talati, Vaskar De, Mayur Tikmani, Reliance Industries Limited
33. Application and Benefits of CombustionONE Solution to Fired Heaters
Yokogawa
34. Innovative Design of FCC Regen Flue Gas Catalyst Sampler
Kavyasree K.V.S.K, Vedula Raghu Kumar, Binoy Das, Bhaskarjyoti Baruah, A.T. Naidu and
B. Balagangadharam
Hindustan Petroleum Corporation Limited – Visakh Refinery
35. Processing Experience of High Acid Crude – Corrosion Control Strategies
A.V.S. Kaushik, Bhaskarjyoti Baruah, A.T. Naidu and B. Balagangadharam
Hindustan Petroleum Corporation Limited – Visakh Refinery
36. Challenges in Procurement of Platinum Based Catalyst
A. Sudhir, K.Vijay Kiran, A.T. Naidu, B. balagangadharam
Hindustan Petroleum Corporation Limited – Visakh Refinery
37. GRM Improvement – Maximizing Diesel & Hydrogen in Refinery with Minimum Capex
B. Gopi, G.V. Madahv, K.Vijay Kiran, A.T. Naidu, B. balagangadharam
Hindustan Petroleum Corporation Limited – Visakh Refinery
38. State-of-the-art methodology for determination of Chloride in hydrocarbon streams
using modified extraction and potentiometric determination
Dr Y S Jhala, Dr Ashutosh Mishra, Dr H K Singh, Dr B R Panda, A K Nath, D Diraviyum,
S Sarkar, D Chakraborty
Indian Oil Corporation limited, Gujarat Refinery
Eni Slurry Technology:
An innovative solution to maximizing value from the Bottom of the Barrel
Nicoletta Panariti #, Raju Chopra *
#
Eni Refining & Marketing
*Haldor Topsoe India Pvt. Ltd.
Introduction
In the next few years, the current trend in both upstream and downstream sectors will continue to require
new technologies which are able to convert heavier and heavier feedstocks into high quality transportation
fuels.
Today a hydrogen addition route is likely to be the right choice due to high conversion, high diesel selectivity
and Euro V grade products. However, conventional hydrocracking solutions like fixed bed and ebullated bed
technologies suffer from limitations on feedstock quality as well as problems related to residue stability that
limits the maximum conversion achievable.
Eni Slurry Technology, the real breakthrough hydrocracking process developed by eni, is the response to the
needs for increased distillate yield and bottom-of-the-barrel-conversion.
EST is a hydrocracking technology featuring:

a very active dispersed (slurry) catalyst, which prevents coke formation and promotes upgrading
reactions (sulphur, nitrogen and metals removal and CCR reduction);

an original process scheme based on a slurry bubble column reactor developed in-house,
perfectly homogeneous and isothermal, that allows optimal control of exothermic hydrocracking
reactions;

a fractionation section for the recovery of the light, middle and heavy distillates;

a system for recycling the catalyst and the partially unconverted fractions.
This process scheme allows the almost complete conversion of the vacuum residue, overcoming the main
limitation of commercially available conversion technologies, i.e. the threshold for the phase separation of
the asphaltenes. Only a small purge is necessary to limit the build-up of metals (Ni and V) present in the
heavy feedstock. Therefore, EST provides higher yields over current available conversion technologies.
Eni Slurry Technology: the Process
From the technological point of view, EST is a hydrocracking process based on the unique features of a
nano-dispersed (slurry) non-ageing catalyst and a special homogeneous isothermal reactor synergistically
working in a novel process scheme that allows an almost total feedstock conversion to distillates.
The active phase of the EST catalyst is unsupported molybdenite (MoS 2) in form of nano-lamellae generated
in situ from oil-soluble precursors. Electron microscopy (HRTEM, High Resolution Transmission Electron
Microscopy) observations reveal excellent dispersion of the catalyst; most of MoS2 is present as single
isolated layers. Stacking phenomena (2 –3 layers particles) involve only a minor part of the catalyst.
Figure 1: High Resolution Transmission Electron Microscopy of nano-lamellae of EST catalyst
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
Since metals precipitate as sulphides forming separated phases without interfering with the naked active
centers of MoS2, catalyst remains practically unchanged during the whole operation eliminating the aging
phenomena. It avoids the need of catalyst substitution (and the relevant plant turn down) typical of all
catalytic hydrotreating processes, Contrary to the conventional supported catalysts utilized in fixed and
ebullating bed reactors, EST catalyst does not suffer for the plugging problems due to the metals and coke
deposits within the pores of the supports. The lower effect of coke, the high surface area, and the absence of
mass transfer diffusion resistances help the catalyst to be more active than supported ones. The very high
specific activity allows therefore the catalyst concentration to be kept at level of few thousand ppm.
Temperature control with a dispersed catalyst is uniform whereas supported catalyst could be subjected to
hot spots. The use of unsupported slurry catalysts is particularly useful in case of feedstock containing high
concentration of pollutants, particularly metals and asphaltenes.
The conversion of heavy products to distillate initiates thermally through breaking of C-C bonds and
generation of free radicals. H-uptake reactions quickly quench them and avoid the chain reaction mechanism
via beta scission of free radical and their recombination that could evolve to coke formation. Indeed the
distance between the MoS2 lamellae in the slurry phase is closer than any supported catalyst to the oil
molecular size, reducing the time elapsed between the radical formation and their hydrogenation on the
catalyst, and contributing to prevent coke growth. The Mo-catalyzed hydrogen uptake allows aromatic ring
hydrogenation, CCR reduction and heteroatoms removal (HDS, HDM and HDN) via the hydrogenolysis of Cheteroatom bonds.
Another important feature of EST is the use of a tailored-designed bubble column reactor operating in slurry
phase. The reactor behaves homogeneously due to the small size of catalyst particles and isothermally due
to the high degree of back mixing fluid-dynamically controlled in the slurry phase ensuring almost flat axial
and radial temperature profiles. It contributes to make the reactor intrinsically safe against temperature
runaway.
The synergetic combination of catalyst development and reactor development enables EST to adopt a
process configuration based on the recycle of unconverted heavy ends achieving overall complete feedstock
conversion, and avoiding fuel oil fatal production of current hydroprocessing technologies.
The following figure shows the simplified reaction section scheme of the EST process.
Figure 2: EST process concept
The heart of the process is a slurry reactor in which the heavy feed is hydrocracked to lighter products in the
presence of the slurry molybdenum based catalyst.
The upgraded oil moves from the reactor to a separation system to recover gas, naphtha, middle distillates
and vacuum gasoil. The gas phase goes to the amine wash section and the clean gas, after recompression
and addition with the make-up hydrogen, is recycled to the reaction section. Distillates are recovered from
the liquid phase, and the unconverted bottom material, together with the dispersed catalyst, is recycled back
to the reactor. Depending upon the feedstock, the process severity (reaction time and temperature) are
optimized in order to generate a residue well within the limit of stability avoiding the phenomenon of
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
asphaltene precipitation that could generate coke and foul the process equipment. The operation of recycling
and blending the partially converted residue with the fresh feed, contributes to maintain stability of the
recycle stream so that it can be reprocessed to get almost total conversion. After repeated cycles, the
system reaches a sort of steady state so that the net result is the total conversion of the feed to valuable
products. A small purge is necessary to limit the build-up of metals (Ni and V) fed with the heavy feed. Purge
is processed to recover residual hydrocarbons and metals including molybdenum. In this way, EST can
handle heavy feedstock assuring very high conversion to distillates because does not generate significant
amounts of by-product, such as coke or heavy fuel oil.
Intermediate products from the slurry section are processed in VGO and Diesel upgrader reactors. These
reactors are installed with special graded bed technology, bulk NiMo catalyst with high HDN/HDA activitiy,
high stability along with high performance reactor internals in order to handle the difficult feed. The reactor
effluents are separated in high pressure loop and finally routed to fractionation sections. Finished products
are separated in fractionation section.
EST Development Road
The original idea of developing in Eni a hydrotreating process based on micronic catalyst goes back to the
late 1980s. After an intensive R&D activity carried out at laboratory level during the 1990s, all the process
steps have been integrated in a 0.3 BPSD Pilot Plant started-up in 2000. Pilot plant operation, mock-up
studies with mimic fluids, and development of suitable models made available all information needed for
designing and constructing a semi-scale 1,200 BPSD Commercial Demonstration Plant (CDP) inside the
battery limits of the Eni Refinery in Taranto (Italy). Since its start-up at the end of 2005, more than 1.000.000
bbl of heavy feed were processed successfully in the CDP. The CDP operation has allowed to consolidate
the know-how on the technology confirming the expected process performance obtained at the pilot scale
and assessing the fluid dynamics of the in-house developed and designed slurry bubble column reactors and
relevant internals. In CDP, Upgrader reactor was loaded with Topsoe NiMo catalyst and equipped with
reactor internals.
One of the main characteristics of the EST process is the excellent feedstock flexibility.
40
35
Maya
C5-Asphaltene wt.%
30
Arabian Heavy
Gorgoglione
Rospo Mare
Visbroken tar
25
Borealis (Oil sands
Bitumen)
20
Basrah
Oil Sands Bitumen
Ural
15
Domestic Residue
10
Russian Atm. Res.
5
0
0
1
2
3
4
5
6
7
8
9
10
Sulphur wt.%
Figure 3: Feedstocks tested in EST Demonstration Unit
With all feedstock used in the CDP, EST has demonstrated the possibility to get the total conversion of the
residue to light, medium and heavy distillates with minimum purge. Moreover, in all cases the process
assures a complete metal removal, an excellent CCR and sulfur reduction and a fairly good denitrogenation.
During the various runs, Topsoe catalysts in Upgrader allowed in all cases to meet the desired product
specifications also in presence of very refractive feeds from EST section; both in VGO and diesel service,
catalysts were quite stable and deactivation rates were found quite low.
The EST/CDP runs have been crucial for developing and consolidating the technology at a size that could
safely allow the scale up to a full-scale commercial plant. Additionally the CDP experience has allowed:
- to learn how to tailor the technology with different feedstock;
- to develop and tune process simulation models;
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
-
to develop and test operating procedures for start-up, steady-state operation and emergency
conditions;
to train operation and maintenance personnel;
to train process engineers;
to evaluate the performance of selected construction materials against corrosion in harsh
environment and;
to evaluate the performance of different kind of instrumentation with heavy, fouling fluids.
The positive results obtained from the Unit encouraged the decision for the first industrial application. Eni’s
Sannazzaro refinery has been chosen to host the first full-scale industrial plant based on EST technology.
The first industrial application: Sannazzaro EST Complex
The first industrial application of Eni Slurry Technology is the EST Complex in eni’s Sannazzaro de’ Burgondi
refinery. The EST Unit has a design capacity of 23,000 BPSD and allows the Sannazzaro refinery to convert
the bottom-of-the-barrel into Euro V diesel and other valuable refinery streams (LPG, naphtha, jet fuel,…).
The plant also represents the first full-scale industrial plant in operation based on a slurry hydrocracking
process.
EST configuration incorporates the most advanced technical solutions and all the operating experience
achieved in more than eight years of continuous tests and operations in the Demonstration Plant at eni’s
Taranto refinery. Many innovations have been brought to the Project, from special items to the up-to-date
construction methodologies which have made extensive use of the pre-assembly of large structures,
foundations and even process heaters.
Sannazzaro EST Complex
• EST unit
- Slurry reaction and separation
-
VGO upgrader (licensed by Haldor Topsoe)
light distillates upgrader (licensed by Haldor Topsoe)
product fractioning
hot oil circuit
effluent treatment
-
pre-reformer and reformer
PSA
fumes DeNOx
-
Thermal sulphur recovery reaction
Catalytic sulphur recovery reaction
Tail gas reduction reaction
Tail gas washing column
Tail gas oxidation
• Hydrogen production plant (licensed by Haldor Topsoe)
- feed desulphurisation and hydrogenation
• SWS unit, sulphur recovery and tail gas treatment
- Double stage SWS
• Amine regeneration unit
• Demi water production and condensate treatment
• Miscellaneous auxiliary systems (cooling water, fuel gas, torch, etc.)
EST Unit successfully started-up on October 2013 and produces high quality distillates, obtained by the
optimize integration and synergism between EST hydrocracking technology and Haldor Topsoe products
upgrading technology.
Main results confirm the proper design of the plant:
- reached 100% design capacity
- slurry reactors perfectly homogeneous and isothermal
- residue conversion ≥95%
- high efficiency of gas-liquid separation (no foaming occurrence)
- product yields and quality as expected
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
Naphtha product
- Specific Gravity 0.707
- Sulfur <3 wtppm
- Nitrogen <3 wtppm
Diesel product (Euro V)
- Specific Gravity 0.840
- Sulfur < 5 wtppm
- Nitrogen <5 wtppm
- Cetane Index (ASTM D-4737) 50
- Polyaromatics < 2.0 wt%
VGO product (marine bunker or HCK/FCC Feedstock)
- Specific Gravity 0.917
- Sulfur <500 wtppm
- Nitrogen <500 wtppm
- Metals <1 wtppm
Figure 4: EST Sannazzaro Complex block flow diagram
The performances achieved in the first three years of operation matched and in some cases exceeded the
design data, in terms of: unit throughput, residue conversion, product yields and properties, hydrogen
consumption and catalyst addition rate.
Moreover, the results demonstrated that the slurry reactors are perfectly homogeneous and isothermal, both
in axial and in radial profiles.
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
Figure 5: Slurry Reactors: Typical Axial and Radial Temperature Profile
The unit is currently in shut down, in order to repair the damages causes by a fire occurred on December
2016. The cause of the event was a mechanical failure, that causes a leakage in a high pressure line. No
injuries to people were registered. The analyses performed by the competent local authorities did not show
unusual concentration of polluting substances in the air.
Eni will take the opportunity of this maintenance activities to incorporate the improvements consolidated in
the first three years of successful operation.
EST potential application:
The EST process has considerable market potential for both upstream and downstream applications. Thanks
to the extremely high feed flexibility, EST may represent the solution for the profitable exploitation of the
huge reserves of unconventional oils, ensuring the availability of additional strategic reserves. In addition,
EST can also be considered an option for the valorization of natural gas reserves, which could be properly
utilized to produce hydrogen for EST process as well as to provide the energy requirements for the process.
On the other side, refining industry could benefit from EST to solve the problem of the bottom-of-the-barrel
upgrading in a very efficient way.
One of the most significant advantages of EST is the possibility of integrating into existing refineries and
petrochemical complexes, creating synergies with existing plants and facilities and minimizing the level of
new investments CAPEX.
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
Conclusions
EST technology allows the almost complete conversion of the vacuum residue, overcoming the main
limitation of commercially available conversion technologies, i.e. the threshold for the phase separation of
the asphaltenes.
EST unit in Sannazzaro refinery has been in successful operation at commercial scale for more than 3 years
with residue conversion more than 92%. Haldor Topsoe VGO and diesel upgraders have been constantly
producing high quality naphtha, diesel and heavy distillates. The highly active and robust catalytic system as
well as the high performance reactor internals has allowed to exceed the expected product quality and cycle
length.
Eni and Haldor Topsoe are in a unique position to offer a full integrated solution to refiners to fully convert all
kind of residue feedstocks i.e. vacuum residue, visbreaker tar, pitch, clarified oil, pyrolysis oil etc. in to all
different kind of high quality distillate products. EST is a robust technology and will enable refiners to fulfil all
their wish list by maximizing value from bottom of the barrel.
EST is delivered by an operator to an operator: the key to success of Eni as EST licensor is the knowledge
achieved from Eni’s experience as the owner and operator of the first EST industrial plant. EST technology
has been developed and fine-tuned after many years of experience in operating the demonstration unit and,
more recently, the industrial plant. This gives clients access to the latest developments and best practices
that have taken time to evolve, thus offering a competitive advantage compared to other technologies
available in the market.
N.Panariti, Raju Chopra: “Eni Slurry Technology: An innovative solution to maximizing value from the Bottom of the Barrell”, 21th RTM
Bottom Of Barrel Upgrading-Future Refinery Configuration
With passage of time, with conditions like global warming and high particulate emission
taking a front seat on the list of global environmental concerns, refinery configurations have
evolved from simple to complex. Over the years, the average crude diet has increasingly
become dirtier due to lesser availability and high prices of light crudes. Thus special
attention is required right from the conceptual stages of setting up a refinery for choosing
the processes and their technologies. Cost effective treatment and conversion processes are
needed to get the best of the products meeting environmental specifications. However
substantial investment is needed in setting up bottom upgradation. The recent downturn in
crude oil prices the incentive to upgrade residues has also shifted as upgrading margins have
been compressed.
Following factors are taken into consideration while short listing the refinery configuration
in a futuristic scenario.







Demand-supply & pricing scenario: current, historical and forecasted
Technologies to meet changing product quality & environmental regulations
Value Maximization from Vacuum Gas Oil (VGO) & Vacuum Residue (VR) Processing
Flexibility of operation
Commercially proven technologies
Capability of capturing future opportunities
Availability & Sourcing of Crude
Shift in Demand-Supply
Product demand and supply play a governing role in selecting a refinery configuration. A
refinery configuration must take into account the type of products that will be marketed as
per the demand supply gap in a particular market segment. Since gasoline and diesel makes
a majority chunk in the overall product portfolio from a fuels refinery, it is logical to see the
trend of gasoline and diesel demand and supply. As per the latest OPEC report, in the period
of 2014-2040, the global Gasoline & Diesel demands are projected to increase at the annual
rate of 0.6 % and 1.0 % respectively. If we look at the Indian Gasoline & Diesel demands,
these are projected to increase at the annual rate of 7.6 % and 5.6 % respectively, from
2016 – 2025. Also, as per five year plan of GoI (2012-17), the projected growth for middle
distillates and gasoline are presented 54% and 20% respectively.
Product Quality Requirement
Regulatory requirements play a critical role in setting up a refinery. For domestic market,
the Government of India proposed to follow BS VI norms by April 1, 2020. With stringent
specifications being evolved, refiners have to make a balance between commercial and
regulatory requirements through a practicable refinery configuration. It becomes now
imperative for a refiner to go for conversion and treating units like Naphtha
Hydrotreatment, Catalytic Reforming, Isomerization, LPG and Kerosene Treatment, Diesel
Hydrotreatment, Sulphur Recovery, Amine Treatment and Sour water Stripping, etc. in their
configuration. Future refinery should be built with capability to produce products complied
with the latest specifications like BS-VI or equivalent.
VGO Processing
Primarily there are two routes for VGO Processing - Processing in FCC or Processing in
Hydrocracker. FCC would be a correct choice if interest is towards Gasoline production and
further planning is for petrochemicals. If interest lies in balancing Middle Distillate
production, then Hydrocracker (HCU) would be beneficial. It yields high quality products,
with no post treatment requirement. Estimated CAPEX for FCC and its associated units is
almost similar to the Hydrocracker unit, including Hydrogen Unit, in the configuration.
Hydrocracker requires relatively higher OPEX due to Hydrogen consumption while the
products from FCC require further treatments amounting to considerable OPEX. When a
cleaner stream like Unconverted Oil from hydrocracker is fed to FCC, it reduces the OPEX in
post treatments of FCC products. There are two options in hydrocracker – Single stage Vs
Two stage. As per our study, single stage HCU with conversion range of 60-90 shows better
prospects, as its Unconverted Oil (UCO) can be the feed to FCC, resulting in a smaller size
FCC unit in the configuration.
Figure 1- illustrates how a three-ring aromatic
compound with several alkyl side chains
reacts in an FCC unit. Methyl and ethyl groups
will tend to stay attached to the aromatic
compounds. Alkyl side chains with a carbon
number of three or greater will cleave off
close to the aromatic ring. The removed alkyl
side chains will initially become olefins, and
may crack again into smaller components
depending on the length of the chain.
Paraffinic compounds with a carbon number
of five or less tend not to crack. Paraffinic
Figure 1
compounds with a carbon number of six crack
slowly, and paraffinic compounds with a carbon number of seven or more will crack fairly
quickly. FCC cracking will not open the aromatic ring structures. Depending on the number
of short alkyl side chains remaining, the compound in Figure 1 could end up in the FCC light
cycle oil product, but would most likely be produced as part of the heavy cycle oil pool or
the decant oil.
If that same three-ring aromatic compound
was partially saturated in an FCC feed
pretreater before being fed to the FCC unit,
the resulting products would be significantly
different (Figure 2). Notice it takes only four
moles of hydrogen to saturate two of the
three aromatic rings. Saturated ring
structures crack open far more easily in an
FCC unit than aromatic ring structures. While
Figure 2
some of the compounds would follow the
upper path and partially dehydrogenate back
to a two-ring aromatic compound, the majority of these partially saturated ring structures
would follow the lower pathway. The longer alkyl side chains would first be cleaved off as
before. The saturated rings would then crack open, leaving a single-ring aromatic structure
and other paraffinic and iso-paraffinic molecules.
The single-ring aromatic structures
created in this way are almost never
benzene. The structures formed will
usually be toluene and xylenes, and thus
have high octane number. The isoparaffins created will continue to crack
to lighter compounds as dictated by the
carbon number. If that same three-ring
aromatic compound was now fully
saturated in an FCC feed pretreater
before being fed to the FCC unit, the
resultant products would again be
Figure 3
significantly different (Figure 3). It would
take seven moles of hydrogen to fully
saturate the three-ring structure. As before, some of the structures would partially
dehydrogenate back to a single-ring aromatic structure, generating several paraffinic
molecules. The majority of these saturated ring structures would crack open completely,
generating numerous normal paraffins, iso-paraffins, and olefins. Again, the paraffins would
continue to crack to lighter compounds, depending upon their carbon number. Normal and
iso-paraffinic molecules have much lower octane numbers than aromatic compounds with
the same carbon number. By fully saturating the initial compound in the FCC feed
pretreater, the gasoline yield would be reduced in favour of a slightly higher LPG and gas
yield, and the octane number of the gasoline created would be lower.
In general, the FCC cracking of
saturated and aromatic ring
Table 1
structures can be summarized
by Table 1. Cracking of multiring aromatics produces high
cycle oil yields but low gas, LPG
and gasoline yields. The
gasoline
produced
would
consist mainly of paraffins and
olefins. FCC cracking of singlering aromatic compounds
produces low gas, LPG and
cycle oil yields, but a high
gasoline yield. The gasoline
produced is mostly aromatic,
and therefore has a high octane number. FCC cracking of saturated ring structures produces
a low yield of cycle oil, a reasonably high gasoline yield, but a higher yield of gas and LPG
than cracking of single-ring aromatics. The gasoline produced would again consist mainly of
paraffins and olefins, and would therefore have a lower octane number.
From this, it can be concluded that if
the refiner’s objective is maximum
gasoline yield from his FCC unit, the
feed pretreater conditions should be
set to ensure that all the multi-ring
aromatic compounds are saturated
down to single-ring aromatics.
Hydrogen addition beyond this point
would only serve to lower gasoline
yield and gasoline octane. Higher
conversions would, of course, be
achieved, but product value would not
be maximized. If a refiner were trying
to use his FCC unit to produce olefins
for alky feed or MTBE or other
Figure 4
chemical production, obviously a
higher feed hydrogen level would
increase the production of olefins.
The curve shown in Figure 4 illustrates the FCC gasoline yield as a function of the feed
hydrogen content. VGOs from more aromatic crudes are near the bottom left-hand end of
the curve, while VGOs from more paraffinic crudes would be located near the top righthand end of the curve. The trend shows that gasoline yield increases with increasing feed
hydrogen. Gasoline yield will not continue to rise, however, as the feed hydrogen content is
increased. At some point, the gasoline yield will drop off, as the conversion to gas and LPG
continues to increase. Overall, conversion will continue to increase, but the gasoline yield
will drop. This drop off is not a detriment if the refiner is trying to make olefins for alky feed
or petrochemical feedstock with the FCC unit, but it is detrimental if the FCC unit is designed
to achieve the maximum gasoline yield. This yield response is represented by the dashed
portion of the curve, although the exact inflection point is a function of the feedstock and
other variables.
It is observed that a FCC – Hydrocracker combination fetches higher margins (~$2 per bbl)
over conventional FCC - VGOHT combination for 350-400 KBPSD refineries. This appears to
be good combination for future VGO conversion as it provides better flexibility between
Diesel and Gasoline.
VR Processing
When a refinery is designed to process heavy crude, yield improvement becomes an
important parameter to consider. Investment in refinery would be attractive only when
yield maximization is achieved. In principle, there are two solutions to upgrade the residue –
these are carbon rejection and hydrogen addition. A brief comparison of these processes is
outlined in the tables below based on the CCR and metal levels in the residue.
There are several technologies available for residue upgrading broadly characterized as
carbon rejection technologies or hydrogen addition technologies
Technologies based on Carbon Rejection
Technology
Conversion Level
Thermal
Cracking
Low
Delayed Coking
Moderate
Solvent
Deasphalting
Feed Specific
(Low to Moderate)
Residue
Gasification
Moderate
Advantages
Challenges
Low
conversion,
FO
generation
Feed stock flexibility, No Coke disposal issue, post
liquid residue generation treatment of liquid products
Operational difficulty with
fluctuating crude
slate.
Low Capex
Needs outlet for Pitch
utilization
No
Liquid
Residue,
Finding optimum utilization
SynGas/H2/Power
of gaseous products
production
Low Capex & Opex
Technologies based on Hydrogen Addition
Technology
Conversion
Level
Fixed Bed
Low
Hydrotreating
Ebullated Bed Residue Moderate to
Hydrocracking
High
Slurry Phase Residue
Hydrocracking
High
Advantages
Challenges
High quality FCC feed
FO disposal
No Solid waste, LSFO Catalyst Disposal, Higher
production
Capex & Opex
Higher Capex & Opex,
High quality products,
Residue Disposal Issue, Very
No Post Treatment
few commercial references
Each of these above technologies, having its own merits and demerits. Selection of most
economic method of Residue Processing technology
is a very complex task due to involvement of
Environmental and Technical issues, e.g., technology
which generates high sulphur Fuel Oil is not
attractive as the global market for high sulphur
residue is expected to decline in the long term with
environmental regulations becoming increasingly
stringent driven by International Maritime
Organization restrictions. While Delayed coking is
still the most favored due to its provenness and
Figure 5
lower capex, Delayed Coking is simple, robust, and
can handle very high levels of feed contaminants;
roughly 25 to 30% of the residue is rejected as petroleum coke, or pet coke. Amongst
hydrogen addition technologies, Ebullated bed residue hydrocracking is coming up as an
attractive option but needs high Capex and has high Hydrogen consumption. Opex of
different technology is shown in figure 5.
Case Study: - Comparison of carbon rejection and hydrogen addition is done for 170000
bpsd crude refinery(24 % VR generation) with respect to conversion, product yield and
product quality. It is observed that liquid conversion of coker units is 65.9 % against the 84.
2 % for Slurry hydrocraker and post treatment of liquid products are not required in case of
hydrogen addition processed unlike coker. But pitch disposal is the challenge for hydrogen
addition technology. Figure 6 - illustrates yield of coker unit & Slurry hydrocraker unit.
H2S/NH3 – 3.5 wt%
H2S/NH3 – 0.6 wt%
Off-Gases – 4.8 wt%
Off-Gases – 3.8 wt%
C3/C4 LPG – 4.7 wt%
C3/C4 LPG – 3.8 wt%
VDU
ATB
Naphtha – 12.3 wt%
Naphtha – 10.6 wt%
V
T
B
COKER
40000 bpsd
H2 – 5.0 wt%
H2 – 25.9 wt%
SHC
Diesel – 53.1 wt%
Diesel – 27.8 wt%
VGO – 18.8 wt%
HCGO – 27.2 wt%
Residue – 5.0 wt%
Coke
– 25.9 wt%
Figure 6
The threshold crude price is $50/bbl brent, for coker technology & hydrogen addition
technology. Beyond this, hydrogen addition technology having greater IRR & NPV over
carbon addition technology.
SHC/EBHC
SHC/EBHC
Figure 7
There are references available for Ebullated bed hydrocracking technology. Slurry
Hydrocracking is also very promising but there are very few commercial references so far.
Suggested Refinery Configuration and Concluding Remarks
Future refinery configuration for upgrading bottom of barrel is concluded as follow:
However, selecting a refinery configuration is always a tedious task because of involvement
of various factors. Following points may be noted for finalizing a new refinery configuration:






There is no single optimal configuration. Product requirement, quality and Capital
investment is the key.
Partial Conversion, Single Stage Hydrocracker coupled with FCC offers an optimum
balance between Gasoline and Diesel.
Ebullated Bed Residue Hydrocracking gives higher margin at higher Capex.
Petrochemical Refinery rather than a Fuels Refinery is the order of the day.
An existing refinery may offer several integration opportunities.
Integrated petrochemicals should be conceptualized in the beginning, while Future
petrochemicals can be built over the refinery in phases.
Multi- zone Catalytic Cracking (MCC) - A new platform for Direct Cracking of
Crude & Opportunity Feedstock to Light Olefins & BTX
Sukumar Mandal, Manoj Yadav, Nilam Limbasiya, Vinod Rayan, Sampath
Nerivetla, Gopal. Ravichandran, Praveen Chinthala, Asit Kumar Das#
Refining R&D Centre, MAB, Reliance Industries Limited, Jamnagar 361142, Gujarat,
India (#Head Refining R&D Centre, Contact email: [email protected]; Mob 9998215210)
Abstract
MCC is a new process developed for direct cracking of crude along with other distress
streams e.g. coker naphtha, slurry oil etc. and low value methanol/DME in sequential
manner in a single riser to make substantial Propylene (> 30wt %) and Ethylene (>18wt
%) and BTX (15%). This technology combines features of three processes namely
Steam Cracking (SC), high severity Fluid Catalytic Cracking (FCC) and Methanol to
Olefins (MTO) in a single riser platform which is unique and first time in the world, giving
advantage of 100-300 $/ton reduction in cost of production per ton of olefins. In this
article, innovative approach of multi-zone catalytic cracking of diverse feedstock in
single riser and stage of development is presented.
Introduction
Steam cracking is the dominant process for production of lighter olefins (ethylene and
propylene) in petrochemical complexes since 1930s. Recent development of shale gas
in North America is revitalizing the olefins industry due to the possibility of using ethane
as feedstock for ethylene crackers. This has forced many industrial and academic
organizations to investigate various catalytic technologies that can be used to convert
directly crude, resids, condensate, tight oil, unstable gas oil, DAO, VGO, paraffinic and
olefinic naphtha, methanol, DME etc. into high yield of light olefins particularly
propylene. Usually separate units e.g. MTO/MTP and two riser/reactor FCCs are
employed conventionally to process such wide range of feedstock.
Reliance has developed a new Multi-zone Catalytic Cracking (MCC) process [1], which
converts such wide range of feedstock including direct cracking of crude to high value
propylene and ethylene in a single riser. The MCC riser has super-high, high,
intermediate and low severity zones where severity in each zone drops gradually in
different sections of riser from bottom to top. On the other hand, the different feed
streams having different crackability need different reaction severity. Therefore, MCC
utilities the variation of cracking rate of different feedstock and intentionally created
different severity zones in a single riser for sequential co-cracking of such wide range of
feedstock. Multi-zone severity of riser also allows to process non-conventional
feedstock like full crude oils as such, condensate, methanol, olefinic naphtha which are
not processable in Steam cracker. In fact, MCC synergizes cracking of light and heavy
feed and endothermic cracking of hydrocarbons with exothermic cracking of
oxygenates, in order to achieve not only a good heat balance but also produces high
yield of light olefins from feedstock, which cannot be processed in Steam cracking.
Features of MCC Process
•
Feed Flexibility
‒ Olefinic C4 from FCC/DCU - C3/C4 Splitter bottoms
‒ Olefinic naphtha - Light and Heavy Coker Naphtha
‒ Straight run naphtha - C6/C7
‒ MCC C4 and non-aromatic naphtha recycle
‒ CSO
•
Optional feedstock
‒ FCC Light naphtha
‒ Opportunity crude - High TAN, Nitrogen, metals except V
‒ Hydrocracker bottom
‒ Methanol - attractive feedstock due its cheaper & stable price
‒ Customized catalyst formulation for high olefins in product
•
Product Flexibility
‒
Ethylene + Propylene maximization
‒
Ethylene + Propylene+ C4 olefin maximization
‒
Ethylene + Propylene+C4 + Gasoline maximization
‒ Various product objective can be made by varying operating conditions,
catalyst composition and feed stock quality
•
Processing
‒ Sequential multi-zone cracking in one riser
‒ Optimized 4 reactor riser zones
‒ Optimum cracking temperatures, based on feed pre-heat
‒ Select zone variables to drive cracking process towards equilibrium
‒ Opportunistic recycle
•
Heat balance
‒ Synergistic combination of light and heavy feedstock
‒ Utilize exothermic heat of methanol cracking
ACE and FCC Pilot Plant Experiments
The MCC concept is to utilize various low value streams available in the refinery as well
as untreated crude to produce high value added light olefins. Many experiments were
performed in ACE unit (Figure1) and FCC pilot plant (Figure 2) to study cracking pattern
of various refinery streams. All potential streams were first cracked in ACE to get the
yield pattern. The cracking temperature for each stream is decided based on sequence
of feed introduction to riser.
Figure 1: Advance Cracking Evaluation (ACE) Unit Figure 2: 2 bbl. /d FCC Pilot Plant
Cracking principle
The cracking of different feedstock e.g., light cracker naphtha (LCN) and n-hexane, was
studied at different temperatures as shown in Figure 3. It can be seen that n-Hexane is
required to be cracked at about 80oC temperature higher than that of LCN to get same
amount of propylene. Further, cracking of LCN at very high temperature is not desirable
as it makes more dry gas as shown in Figure 4. Therefore, it can be concluded that
each hydrocarbon stream needs to be cracked at optimum severity to get optimum
yield.
N-Hexane
LCN
N-Hexane
Propylene Yield, Wt%
DG-Ethylene Yield, Wt%
LCN
610
620
630
640
650
660
670
680
Temperature, ˚C
Figure 3: Cracking of LCN and n-Hexane at
different temperature
610
620
630
640
650
660
670
680
Temperature, ˚C
Figure 4: Dry Gas selectivity at
different temperature
Cracking Zone suitability for diverse cracking streams
In order to provide optimum cracking severity corresponding to cracking behavior of
different streams, the riser is divided into multi-zones as shown in Table-1. Three to four
hydrocarbon feeds are injected along the entire length of the riser i.e. C 4 hydrocarbon
and/or C4-C6 paraffin in the bottom zone, olefinic naphtha having 5 to 12 carbons in the
middle zone, and heavy hydrocarbon feed and oxygenates in the top zone of the riser.
Temperature in the zones is in the range between 640 to 680ºC, 630 to 650ºC, 600-640
O
C and 500 to 620ºC, respectively, and a weight hourly space velocity (WHSV) in the
zones is in the range of 1 h-1 to 10 h-1, 50 h-1 to 100 h-1,100 h-1 to 150 h-1 and 150-200 h1
, respectively to sequentially crack these hydrocarbon feedstock.
Table-1, Typical operating condition of different riser zones
Zone
Temp 0C
WHSV, /hr
Severity
4
570 – 620
150 – 200
Low
3
600 – 640
100 – 150
Moderate
2
630 – 650
50 – 100
High
1
640 – 680
1 – 10
Super
high
In accordance with the principle explained above, olefinic C4 hydrocarbons stream
along with recycle C4 stream is fed at the riser bottom where temperature is maintained
at 680˚C. Straight run naphtha and recycled lighter naphtha streams are fed above the
C4 streams. Temperature is maintained at 650˚C in this part of riser. In addition to these
streams, light coker naphtha and heavy coker naphtha are fed where riser temperature
is maintained at 640˚C. As coke yield of these lighter streams is very less to sustain
heat balance of the unit, additional coke requirement is fulfilled by cracking of heavier
streams like CSO. Cracking pattern of CSO shows upgradation of this stream and it
also provides majority of coke requirement to maintain heat balance. Riser temperature
at this location is maintained at 620˚C. In case, crude is available, the same can be
injected at third zone. The unconverted streams can be recycle backed to zone1 and
2.Therefore, MCC process can be applied for only untreated crude provided
contaminant like vanadium is in within limit.
Thermodynamic Equilibrium of Light Olefins
The single riser concept with mixed feed has limitation on maximum propylene yield due
to thermodynamic equilibrium [2] at conventional FCC operation. However, in spite of
single riser cracking, MCC C3= yield is close to thermodynamic equilibrium limit (Figure5) of light olefins, due to
•
Higher riser temperature is allowing shift of equilibrium and substantial cracking
paraffinic naphtha
•
Substantial cracking of C4, C5 and higher olefin precursors at the bottom of riser
•
Optimum condition for each zone to maximize propylene and ethylene depending
on crackability of each feed streams.
Optimum catalyst formulation
Selecting catalyst composition for MCC process is also very important aspect as
several stream of different cracking behavior are to be cracked in single riser.
Several experiments were carried out in ACE by varying catalyst composition
comprising medium sieve zeolite such as ZSM-5 and large pore zeolite e.g., Y
zeolite. RIL developed ZSM-5 additive and metal modified ZSM-5[3, 4]. RIL’s ZSM 5
additive gives higher yield of light olefins i.e., propylene by 1.2 wt% and ethylene by
1.3 wt% in comparison with commercial ZSM-5 additives at similar condition.
Moreover, metal doped ZSM-5 additive produces better propylene yields
comparable to commercial additive, but additionally reduces yield of ethylene(1.9
wt%) & dry gas( by 1.5 wt%). This metal modified additive is useful when higher
ethylene and dry gas yield are not desirable due to site specific constraints.
MCC vs other catalytic cracking technologies
Recently extensive efforts are being made to maximize light olefin production from
refinery complex to meet high demand of propylene and ethylene. As a result,
different enhanced FCC processes were developed and some of them are
operational. For example, Maxofin, PetroFCC, MILOS, DCC, Indmax etc processes
are employed to crack heavy hydrocarbons (hydro treated VGO, VGO, HVGO, RCO
etc) at very high reaction severity using admixture of FCC catalyst and ZSM-5
additive or single catalyst formulation containing both Y and ZSM-5 zeolites. Some
other processes like Superflex, ACO, PCC etc. are developed to convert low value
naphtha feedstock to light olefins by employing ZSM-5 zeolited based catalyst at
very high reaction severity in FCC like circulating fluid bed configuration. As shown
in Figure 6, MCC is capable to crack light to heavy feedstock in single riser.
Reaction Severity (ROT)
700
650
ACO/Superfl ex
CPP/S&W
MCC - C4 to Resid/Crude + Oxygenate in Single Riser
600
RxPRO/PCC
DCC + Bed crac
RxPRO/PCC
PetroFC C
550
Gaso line FCC
500
VGO/Resid
Naphtha
0
10
20
30
40
50
Feed Carbon Number
Figure 6, Different Process for Max light olefin generation
MCC Applications in Existing Refinery and Petrochemical complex.
In existing refinery, MCC can be integrated in following scenario:
•
•
•
•
Upgrading low value refinery streams e.g. Light coker naphtha, Heavy coker
naphtha, Visbracker Naphtha, Condensate, CSO, Resid to high value
Petchem feedstock to Light olefins, BTX, Heavy Aromatics.
Crude to Olefins – Direct cracking of crude wherein no CDU/VDU/Flashing is
required. Also it can be combined cracking of condensate, shale oil, tight oil
etc. along with Crude. About 120 crudes across world have been scanned as
suitable for direct processing in MCC
Gasoline Quality – Limited cracking of Gasoline streams to reduce olefin
content while increasing light olefins production
Synergy with other Bottom Upgradation Projects e.g. Resid HC and SDA by
processing heavier streams produced from these units to light olefins
MCC process can be integrated with steam cracking for light olefin maximization
•
MCC integration with Steam Cracker (SC) – Cracking of C4, C5-C8 olefinic
raffinates from steam cracker and Py tar in MCC riser while SC can crack C3 and
C4 paraffins from MCC. In case, ethane is used as feed in SC, SC propylene
production drops, which can be enhanced easily by adopting MCC.
•
Methanol / DME cracking in MCC riser & Integrate to SC
MCC concept commercialization in SEZ FCC
In order to prove this concept, FCC pilot plant was operated with VGO to generate
base line data and then with LCN Co-cracking with VGO as shown in Table-3.
In Reliance SEZ FCC unit, there are four light coker naphtha (LCN) injection
nozzles at few meter below the VGO feed injection nozzles. Two plant trials were
carried out as shown in Table-2. It can be seen that LCN produces propylene in the
range of 20 to 23 wt% in plant vs. 19 wt% in pilot plant. The marginally lower
propylene predicted from pilot plant is due to mainly relatively lower severity.
Another, important point is to mention that LCN co-cracking makes lower LCO and
CSO yield. This is because catalyst to oil ratio goes up for satisfying heat demand in
riser for supplying heat for LCN cracking and vaporization. This pilot plant data and
commercial trials proved the concept of two feed co-cracking in single riser.
Table 2: Pilot Plant data vs. Plant data: Two feed Co-cracking of LCN & VGO
Col-1
8.57
23.0
-23.57
-2.86
Pilot with heat
balanced
Col-3
7.08
18.66
-25.77
-0.91
B
B+0.05
B
B
B+0.30
B-0.30
Plant trial
Yield, wt%
DG,
Propylene
LCO
CSO
Process Condition
Riser top pressure, kg/cm2(g)
C/O ratio
WHSV, hr-1
MCC economic
Preliminary economic of MCC process was calculated and compared with steam
cracking as shown in Table 3.
Table 3, Comparison of MCC economic with Steam Cracker
Product
LCN
Crude
Lt SRN
MCC+
MCC+
SC
Ethylene
24.8
20.5
34.0
Propylene
BTX
Life cycle cost, $/T of
E&P
Payback, yrs
40.5
18.5
33.0
11.5
18.0
12.7
B-100~150
B-200~300
B
C-0.5~0.8
C-1~1.5
C
Superior economics in MCC over SC is mainly due to following benefits
•
Major benefit in MCC arises from use of low value streams as feedstock which
SC can’t handle. This spread over naphtha translates into lower cost of
production
•
MCC handles olefinic feed directly whereas SC requires hydro treating
•
Direct crude cracked in MCC + recycling of naphtha streams in bottom zone of
riser, producing higher yield of light olefins
•
Lower dry gas yield in MCC vs SC
•
Good heat balance in MCC, coke from heavy &
methanol,
•
Large scale MCC plant vs multiple furnace in SC => lower capex
exothermic cracking of
Summary
MCC is a new process developed for direct cracking of crude along with other
distress streams e.g. coker naphtha, slurry oil etc. and low value methanol/DME in
sequential manner in a single riser to make substantial Propylene (> 30wt %) and
Ethylene (>18wt %) and BTX (15%). This technology combines features of three
processes namely Steam Cracking (SC), high severity Fluid Catalytic Cracking
(FCC) and Methanol to Olefins (MTO) in a single riser platform which is unique and
first time in the world, giving advantage of 100-300 $/ton reduction in cost of
production per ton of olefins.
References
1. Sukumar Mandal, Manoj Yadav, Amitkumar Parekh, Asit kumar Das,
Shubhangi Jaguste, Praveen Kumar Chinthala, Gopal Ravichandran, Mahesh
Marve, Ajit Sapre, US Patent no. 9550708
2. Xiaoping Tang Æ Huaqun Zhou Æ Weizhong Qian Æ Dezheng Wang Æ
Yong Jin Æ Fei Wei, High Selectivity Production of Propylene from n-Butene:
Thermodynamic and Experimental Study Using a Shape Selective Zeolite
Catalyst, Catal Lett (2008) 125:380–385 DOI 10.1007/s10562-008-9564-8
3. Srikanta et al. INP 275804, CN2072483, US20140116923A1, “Process and
Composition of a New Catalyst/Additive for Reducing Fuel Gas Yield in Fluid
Catalytic Cracking (FCC) Process”.
4. Ravichandran et al. US patent no. 9067196, “FCC catalyst additive and
preparation method thereof”.
Progress and Future of HDPE Technology
Shin-ichi Kojoh, Ph.D.
Director of License Department and Director of Technology Department.
Licensing Division, Mitsui Chemicals, Inc.,
Shiodome City Center, 1-5-2, Higashi-Shimbashi, Minato-ku, Tokyo 105-7122, Japan
[email protected]
ABSTRACT
CX process is bimodal slurry process for producing high density polyethylene
(HDPE) developed by Mitsui Chemicals, Inc., having high reputation in quality of
products. It has been licensed to 18 countries and total capacity of CX licensees
has reached over 7,000,000 tons per year. Not only process development but also
development of MgCl2-supported TiCl4 catalyst brought about “Process Innovation”
and “Product Innovation” in HDPE industry. Supporting TiCl4 on MgCl2 gives two
important effects on Ti active sites. One is electron donation from MgCl 2, resulting
in extremely high catalyst activity. Another is active site location only on surface,
contributing to purified nature of active site. It enables only one catalyst system to
produce a wide variety of grades in CX process.
The progress for establishing CX process and its benefits are introduced
together with its future direction such as use of metallocene catalyst.
INTRODUCTION
Prof. Ziegler’s discovery of TiCl3 catalyst system for polymerizing ethylene initiated
HDPE industry in 1950’s. However, its catalyst activity was so low that HDPE producers
had to extract catalyst residue form the resulting polymer with alcohol. Therefore, HDPE
production process needed a lot of energy and production cost due to the catalyst removal
and solvent recovery. An enormous effort has been devoted into improvement of the
catalyst. Eventually, invention of MgCl2-supported TiCl4 catalyst system brought about
process innovation and product innovation. Not only drastic decrease of HDPE production
cost but also precise polymer-structure control was realized by this catalyst system.
Furthermore, combination of process development and this catalyst system enabled us to
produce high quality product in a wide variety of HDPE applications such as film, pipe,
bottle, container, drum and automobile fuel tank. Thus, HDPE industry has grown rapidly by
replacing such materials as paper, wood or metal and the world-wide capacity of HDPE has
reached over 40,000,000 tons per year. This paper introduces how and why such
innovation was realized and what future direction of HDPE industry is.
CATALYST INNOVATION
It was in 1953 that Prof. Ziegler discovered the TiCl3 catalyst system for polymerizing
ethylene. Then, in 1955, Mitsui Petrochemical Industry Ltd (current Mitsui Chemicals, Inc.)
obtained patent license directly from Prof. Ziegler and carried out its scale up to commercial
size by themselves. Thus, Mitsui Petrochemical Industry Ltd became the first HDPE
producer in Japan in 1958. However, catalyst activity of the TiCl3 catalyst system was so
low that the resulting polymer was colored by high content of catalyst residues, which also
caused rust of molding machines. Therefore, HDPE commercial production process needed
a step to remove the catalyst residues with alcohol. Besides, alcohol deactivates the
catalyst and should have been careful about recycling solvent after using alcohol.
It was sincerely desired to increase the catalyst activity and a lot of effort was devoted.
There were two directions to increase the catalyst activity at that time. One was to increase
specific surface area of TiCl 3 and another was to support Ti active species on surface of
inorganic materials. The former could improve the catalyst activity but there was limitation
to bring about catalyst innovation by modifying TiCl3. The latter was a difficult way, because
it was hard to support Ti active species on such conventional support as silica or alumina.
Eventually, catalyst innovation was realized by supporting Ti active species on MgCl2,
leading to more than 100 times higher activity than the TiCl3 catalyst [1]. Among a lot of
inventors applying patents about MgCl2-supported TiCl4 catalyst system, Mitsui
Petrochemical Industry Ltd is the first patent applicant in the world about this catalyst
system [2]. Reasons why MgCl2 was so successfully used as a support are considered that
both crystals of MgCl2 and TiCl3 belong to the same hexagonal modification and Mg 2+ and
Ti4+ have similar ionic radii, 0.68 and 0.65 Å, respectively [3].
Electron donation from MgCl 2 to Ti active species is considered to contribute to such
extremely high activity. Because electron negativity of Mg 2+ is smaller than that of Ti 3+,
electron is donated from MgCl2 to Ti active species. It was proposed that such electron
donation brings about back-donation of d-electrons of Ti to π*-orbital of coordinated
monomer, whileπ-electrons of the monomer are donated to a vacant d-orbital in Ti [4]. It
promotes the monomer to be inserted into growing polymer chain
chain-propagation rate constant (kp) in polymerization reaction is boosted [5].
and
the
PROGRESS IN HDPE PROCESS AND PRPDUCT
The extreme enhancement of the catalyst activity enabled us to omit the step to
remove catalyst residues from HDPE production process. It enabled us to simplify the
process and enlarge the plant scales. Thus, a lot of energy and production cost could be
cut and the process was established as CX process. CX process has been licensed
worldwide since 1970’s.
Figure 1. Licensing experiences of CX process
As shown in Figure 1, its maximum plant size was enlarged from 60,000 tons per year in
1970’s to current 500,000 tons per year and total capacity of its licensees reached over
7,000,000 tons per year.
Supporting Ti active species on MgCl 2 not only increased the catalyst activity but also
purified Ti active site nature. Because Ti atoms, which could become Ti active species by
reduction reaction with alkyl aluminum, existed not only on surface but also inside of crystal
of TiCl3 catalyst, active site nature of the TiCl3 catalyst was very heterogeneous, reflected
by its various Ti’s locations. Supporting Ti active species only on surface of MgCl2 made
active site nature more purified and enabled us to control polymer structure precisely. Ideal
polymer structure for high value added HDPE is shown in Figure 2, consisting of high
molecular weight portion and low molecular weight potion.
Figure 2. Concept of ideal polymer design
Low molecular weight portion contributes to high processability of HDPE and high
molecular weight potion brings about excellent mechanical strength, which is enhanced by
comonomer incorporation for short chain branches.
CX process is bimodal slurry process designed for making molecular weight
distribution broader and comonomer distribution narrower than other processes. Besides,
CX process has a step to remove low-tail polymer (LP) in its molecular weight distribution,
leading to much better properties of HDPE product. Furthermore, LP is useful as wax in
such application as modifier for polyvinyl chloride. Therefore, CX process is the most
suitable to produce such bimodal HDPE as film, large-blow or pipe grades.
FUTURE DIRECTION
Owing to suitable design for the purified Ti active site nature of MgCl 2-suppoted TiCl4
catalyst system, CX process can provide a wide variety of HDPE grades in unimodal or
bimodal only with one catalyst system, although other processes need more catalyst
systems to cover all the HDPE grades. Such CX process design has been applied for
metallocene catalyst system. It realizes narrower molecular weight distribution, narrower
comonomer distribution and higher flexibility of selecting amount and kind of comonomer.
Thus, further purification of active site nature from MgCl2-suppoted TiCl4 catalyst
system to metallocene catalyst system is expected to realize further product innovation. For
examples, high rigidity, impact resistance, long-term resistance and chemical resistance are
expected, contributing to resource saving such as lightening container's weight or reducing
film thickness [6]. Then, it enables lower molecular weight polymers to maintain enough
mechanical strength, leading to energy saving by increasing molding speed.
Conclusively,
combination
between
purified
nature
of
catalyst
and
high-quality-oriented process design realizes ideal polymer structure, letting HDPE
producers to enjoy differentiated commodity market.
References
[1] Kashiwa, N. J Polym Sci Part A: Polym Chem 2004, 42, 1.
[2] Kashiwa, N.; Fujimura, H.; Tokuzumi, Y. JP Patent 1031698 (application in Japan Aug. 1,
1968).
[3] Kashiwa, N. Polym J 1980, 12, 603.
[4] Doi, Y.; Soga, K.; Murata, M.; Suzuki, E.; Ono, Y.; Keii, T. Polym Commun 1983, 24,
244.
[5] Kashiwa, N.; Kawasaki, M.; Yoshitake, J. In Catalytic Polymerization of Olefins, Keii, T.;
Soga, K., Eds.; Kodansha-Elsevier, Tokyo, 1986, p. 43.
[6] http://www.primepolymer.co.jp/english/product/pe/evolue-h.html
2nd generation HyBRIMTM catalyst: Innovation to fulfill the refiner’s wish
Authors: Michael Tinning T. Schmidt, Raju Chopra and Hussain Lokhandwala
Organization: Haldor Topsoe A/S
Every refiner around the world has a big wish – increase the profit of operation! All Indian refineries
are targeting to meet their commitment of producing 100% BS IV diesel fuel specifications by Q1
2017 and 100% BS VI by year 2020. Addressing the increasing demand for cleaner products and
more volume swell – while still ensuring cost-efficient operations – is a very challenging task. Hence,
continuous improvement in the activity and stability of the hydrotreating catalysts is imperative,
particularly when the implementation of BS-VI standard in India is just around the corner.
A competitive hydrogen production cost makes it very profitable to catalytically add hydrogen to
middle distillate fractions, thereby increasing the liquid volume swell, resulting in higher yields of
valuable diesel. To implement this strategy while balancing the costs, the refineries need better cost
efficient alumina-based catalysts with highest possible hydrogenation activity in order to achieve
desired boost in unit performance. Furthermore, alumina-based hydrotreating catalysts will help
minimize the operating cost when targeting volume swell in comparison with higher cost unsupported
catalyst formulations.
Topsoe has always been on the forefront of development in both hydrotreating catalysts and
technology. With its in-depth knowledge about catalysis and advanced R&D facilities, Topsoe
developed and launched HyBRIMTM in 2013 that involves an improved production technique for both
NiMo and CoNiMo hydrotreating catalysts. It combines the BRIM® technology with a proprietary
catalyst preparation step. Better understanding of the interaction between active metal sites and the
alumina carrier has now enabled Topsoe to launch the 2nd generation HyBRIMTM catalyst in 2016:
TK-611 HyBRIMTM. This launch comes in just 3 ½ years after introducing TK-609 HyBRIM™ which
took the industry by storm. Today TK-609 HyBRIM™ is installed successfully in more than 120
hydroprocessing units. With the 2nd generation TK-611 HyBRIMTM, Topsoe has been able to boost
the activity by 25% compared to the 1st generation TK-609 HyBRIMTM catalyst without compromising
on the catalyst stability. This corresponds to 5-6 °C lower start-of-run temperature at otherwise
unchanged conditions.
This increased activity combined with same stability provides additional flexibility to refiners to obtain
longer catalyst cycles, more throughput, better product qualities or processing of more cracked
feedstock, which can be utilized to improve the overall profitability and economy of all
hydrocracker/FCC pretreating and ULSD units. The focus of this paper will be on our latest NiMo
innovation TK-611 HyBRIM™, which is part of the 2nd generation HyBRIM™ technology. It describes
the improvement in Topsoe HyBRIM™ technology and demonstrates the catalytic possibilities with
TK-611 HyBRIM™ and will show its superiority with respect to activity as well as stability in
comparison with TK-609 HyBRIMTM catalyst.
Continuously improving alumina-based catalysts
Topsoe has been at the core of the most important developments in the hydroprocessing catalyst
industry for over 4 decades and is today the market leader in ULSD and pressure drop control. In the
1970s, Topsoe discovered the active CoMoS phase in hydrotreating catalyst which revolutionized
the catalyst world by applying fundamental research in the catalyst development approach. In the
1980s, Topsoe researchers, headed by Dr. Henrik Topsoe, discovered the difference between
Type I and Type II hydrotreating catalysts and gave them the names that are known throughout the
industry today. With this breakthrough, the hydroprocessing catalyst development entered the
nanotechnology era, and the Type II hydrotreating catalysts became the industry standard for highactivity catalysts. In the early 2000s, Topsoe’s dedicated commitment to fundamental research in
surface science paid off again, and a new activity site was discovered by Scanning Tunneling
Microscopy (STM) and in-situ Transmission Electron Microscopy (TEM). This layer is located at 1
atom layer from the edge of the metal slab and were named BRIM® sites. Not on the edge/rim of a
hat, but a little further in on the brim of the hat. This is illustrated by the yellow color on the TEM
pictures of a CoMo BRIM and NiMo BRIM sites, shown in Figure 1.
Figure 1. TEM pictures of CoMo BRIM and NiMo BRIM sites
It was discovered that the pi-electrons in the outer orbit of the sulfur atoms, located on the top of the
metal slab one atom layer in the form of the edge, are not as tightly bounded to the sulfur atom as
the pi-electrons around other sulfur atoms on the basal plane. These loosely bounded pi-electrons
basically generate a cloud of pi-electrons, shown in Figure 1 as the grayish glowing circles on the
basal plane, were named brim sites. It was found that brim sites are responsible for the first step in
the hydrogenation route in the removal of the most refractive sulfur compounds in both diesel and
vacuum gas oil fractions. The Type II direct desulfurization sites are located on the edges of the
sulfided CoMo/NiMo slabs. These sites are responsible for the second step in the hydrogenation
route, i.e. the extraction route.
With this finding, Topsoe developed the BRIM® technology – both within the CoMo and NiMo type
catalysts, which fueled Topsoe’s unparalleled growth in market share globally due to a top tier
catalyst portfolio. Over the years, Topsoe has researched and understood the interaction between
the active sites and carrier phase. This development has mainly been possible because of Topsoe’s
improved carrier preparation technology and design of the alumina crystallites.
From BRIM® to HyBRIM™
Topsoe’s latest catalyst technology, HyBRIM™, involves an improved production technique for both
NiMo and CoNiMo hydrotreating catalysts. It combines the BRlM® technology with a unique
proprietary catalyst preparation step. The synergy effect from merging the two technologies has
enabled Topsoe to design an advanced metal slab structure that is characterized by an optimal
interaction between the active metal structures and the catalyst carrier. The activity of the Type II
sites is largely linked to this interaction between the metal slab and the carrier. The HyBRIM™
technology exploits this interaction and substantially increases the activity of both the direct sites and
the hydrogenation sites without compromising the catalyst stability.
Topsoe’s NiMo catalysts that are developed today are around three times more active than the
catalysts produced in the 1990s. To put that in perspective, TK-611 HyBRIMTM has the same metal
content as TK-555 but is 20°C more active!
Figure 2 below illustrates the development of our many catalyst generations. Since Topsoe’s
scientists employed sophisticated tools, such as electron microscopes, in situ monitoring, and highthroughput screening in their research programs, we have made considerable progress within the
hydrotreating and hydrocracking catalyst development. The BRIM® and HyBRIM™ catalyst
technologies are the direct outcomes of this scientific approach.
Figure 2 Topsoe’s catalyst development progress – from Type I through 2
nd
generation HyBRIM™.
The novel HyBRIM™ technology was originally introduced with Topsoe’s TK-609 HyBRIM™ in 2013
and has since been extended to include several different hydrotreating catalysts covering mediumto high-pressure refinery applications. As mentioned before, the need for ULSD production from lowquality crudes and higher severity hydrocracking is creating a need for even better catalysts.
During the past three years, the HyBRIM™ technology has been broadly recognized by the industry
to be at the forefront of hydrotreating. In addition, more than 120 hydrotreating units around the
globe have a HyBRIM™ catalyst installed right now. However, Topsoe’s researchers recently
discovered even more potential within the HyBRIM™ method – a potential of utilizing the active
metals to an even higher extent and securing a dispersion of active sites to a level never seen
before. Therefore, Topsoe is extremely proud of being able to launch the 2 nd generation of the
HyBRIM™ technology – the TK-611 HyBRIM™ catalyst with 25% higher activity for both sulfur and
nitrogen removal.
Figure 3 Image of TK-611 HyBRIM™
Topsoe has found that not only the dispersion of metals is important, but the interaction between the
metal slabs and the carrier material also plays a very important part. Topsoe was the first to discover
there are two different states of the active sites for desulfurization. And very originally, chose to
called these two type of sites for Type I and Type II sites. R&D found that the activity increases with
increasing amount of active sites. And what was even more important is that the activity of Type II
sites increases significantly more than the activity of the Type I sites. It was seen that the Type I
sites has a stronger interaction with the carrier.
Therefore, in order to have more activity, more type II sites are needed which have reduced
interaction with the carrier material. One way of reducing interaction could be to stack the
CoMoS/NiMoS slabs or increase the promotor (Co/Ni). However, stacking or clustering is at the
expense of reduced metal dispersion. Just adding Co/Ni-sulfides do not increase the activity since
the HDS/HDN activity follows CoMoS/NiMoS phase. At Topsoe we have found a way to reduce the
interaction with the carrier for single slabs, so multiple metal slabs can be dispersed into single
slabs, but still with Type II sites, thereby getting increased number of active sites, using the same
amount of metals.
Topsoe succeeded in producing a catalyst by creating BRIM sites on the top slab and type II sites on
the edge without stacking the metals slabs. This innovative improvement resulted from an improved
metals impregnation and decomposition technology, along with improved alumina carrier design and
preparation.
Increased activity, same stability
When using the latest generation TK-611 HyBRIM™ catalyst in either ultra-low sulfur diesel (ULSD)
or in a VGO hydrocracker pretreatment service, the advantages for the refiner are many.
Obviously, a higher activity means lower SOR temperatures for the same feed and capacity
compared to previous cycles, which implies longer cycle lengths. Or, the new high activity can also
be utilized in terms of increasing unit throughput for the same cycle length. In addition, the TK-611
HyBRIM™ will also increase the volume swell due to better hydrogenation functionality in terms of
HDS, HDN and HDA. In the case of a hydrocracker, TK-611 HyBRIM™ will give a lower nitrogen slip
from the pretreat section to the cracking section, resulting in higher conversion and better yields.
Refiners can benefit from purchasing more opportunity crudes or processing more LCO and with TK611 HyBRIM™. In any case, the TK-611 HyBRIM™ will significantly improve the profitability of the
refinery assets.
Figure 4 shows the comparison between TK-611 HyBRIM™ with TK-609 HyBRIM™ side-by-side. It
is seen that if the two catalysts are operated for the same product sulfur of 10 ppmw, TK-611
requires 6°C lower temperature. In other words, at the reactor temperature of 328°C for getting 10
ppmw product sulfur with TK-611 HyBRIM™, the TK-609 HyBRIM™ catalyst delivers a product with
32 ppmw sulfur. This is a remarkable step-change in activity for ULSD.
Figure 4 Comparing TK-611 HyBRIM™ with TK-609 HyBRIM™ in ultra-low sulfur diesel service
The same type of experiment is illustrated in Figure 5 for vacuum gasoil (VGO) at hydrocracker
pretreatment conditions. While 62 ppmw product nitrogen slip is achieved with TK-609 HyBRIM™,
the TK-611 HyBRIM™, at exactly the same conditions, is able to deliver a product nitrogen slip of
only 26 ppmw. It can also be seen that VGO product sulfur of 193 ppmw is achieved with TK-611
HyBRIM™ as compared to 322 ppmw as achieved with TK-609 HyBRIM™ for the same feed at the
same conditions. Such difference is a substantial improvement for a hydrocracker or a FCC
pretreating unit.
Figure 5 Comparing TK-611 HyBRIM™ with TK-609 HyBRIM™ in VGO service
A high start-of-run activity has no real meaning unless it is accompanied with high catalytic stability,
ensuring that the improved performance is maintained over the projected cycle.
During the development of the BRIM® and subsequent HyBRIM™ technologies, Topsoe’s
researchers have observed how the catalyst preparation steps influence the catalyst functions. This
knowledge has resulted in preparing very active and stable CoMoS/NiMoS catalyst formulations and
alumina carrier phase. The data shown in Figure 6 compares the stability of the new TK-611
HyBRIM™ with TK-609 HyBRIM™ in VGO service at the same operating conditions. The testing
reveals that even though TK-611 HyBRIM™ is operated at higher sulfur and nitrogen conversion
levels, due to its higher activity, the two catalysts exhibit exactly the same performance stability.
Figure 6 Catalyst stability for TK-611 HyBRIM™ vs TK-609 HyBRIM™ in VGO service
Higher volume swell
The lower cost of hydrogen and growing worldwide middle distillate demand make it profitable to
pump hydrogen into the middle distillate fraction in the refinery, thereby increasing the liquid volume
swell and giving the higher yields of valuable diesel product – a very profitable scenario for the
refiner when excess hydrogen is available. The term ‘volume swell’ refers to the increase in the liquid
volume when the product density and distillation are lowered by hydrotreating. This includes
contributions from removing sulfur & nitrogen, hydrogenation of olefins and most importantly,
saturation of aromatics.
Topsoe has previously demonstrated that there is a strong correlation between the observed density
improvement and the degree of aromatic saturation (%HDA). The aromatics saturation is controlled
by the catalyst’s hydrogenation activity, the presence of inhibitors, such as nitrogen species, and by
thermodynamic equilibrium. The overall amount of hydrogen consumed by hydrodearomatization
(HDA) reactions is also dependent on the actual amount of aromatics present in the feed, as shown
in Figure 7.
100%
LGO
Volume swell, vol%
P = 90 bar
LHSV =
0.92 hr-1
T = 321°C
10,0
9,0
8,0
30% LGO
70% LCO
7,0
P = 90 bar
SV = 0.92 hr-1
T = 338°C
30% LGO
70% LCO
6,0
P = 90 bar
SV = 0.92 hr-1
T = 327°C
30% LGO
70% LCO
P = 90 bar
SV = 1.5 hr-1
T = 338°C
5,0
43% LGO
57% LCO
4,0
P = 117 bar
SV = 2.2 hr-1
T = 346°C
100% LGO
3,0
P = 90 bar
SV = 0.92 hr-1
T = 321°C
2,0
1,0
Condition 1
Condition 2
Condition 3
Condition 4
Condition 5
Feed A
Feed B
Feed B
Feed B
Feed C
0,0
Figure 7 Obtained volume swell with HyBRIM™ catalyst in a LCO hydrocracker
A simple way to quantify the volume swell is to compare the liquid product density with the feed
density. This is an easy approach and is based on available data; however, it does not take the yield
losses into account. The correct and more accurate way is therefore to compare the increase in the
C5+ volume yield with the fresh feed.
Topsoe’s researchers have previously published that the presence of nitrogen and, in particular,
basic nitrogen compounds, strongly inhibits the HDA reactions. Therefore, it is beneficial to remove
nitrogen to very low levels, i.e. lower than 2–3 wt ppm N, in order to increase the HDA and the
associated volume swell. Consequently, TK-611 HyBRIM™ with maximum activity for nitrogen
removal (HDN) will also provide the highest volume swell.
Figure 7 illustrates the effect on volume swell when changing the operating conditions and feed
blend. The pilot plant test was conducted simulating a LCO hydrocracker pretreat unit using a
HyBRIM™ catalyst as the main treating catalyst. It included five different conditions, where pressure,
temperature, space velocity, and LCO amount in the feed blend were varied. The results show that
the HyBRIM™ catalyst removed the nitrogen to very low levels, i.e. below 0.2 wt ppm N, which is the
detection limit. Consequently, this indicates that saturation of aromatics and, in particular, monoaromatics should take place in the last part of the reactor. In this pilot plant test, the temperature has
been kept in the low range in order to avoid thermal cracking. As expected, Feed B, being very rich
in aromatics due to the high LCO content, obtained the highest volume swells at test conditions 2, 3,
and 4.
In Figure 8, the corresponding aromatics content from Feed B is plotted. It is seen that highest
volume swell is reached where the most aromatics are saturated, which is condition 3. This can be
explained by the higher reactor temperature and low space velocity. Especially the mono-aromatics
are saturated the most at condition 3, and there is actually a correlation between the obtained
volume swell and the residual content of mono-aromatics. In refinery terms, the highest volume swell
observed with Feed B corresponds to gaining more than 3,300 bpd additional liquid out of a 40,000
bpd unit. This is remarkably high in conventional hydrotreating without any cracking catalyst and is
the result of a high degree of aromatic saturation.
80
70
Aromatic content, wt%
60
Feed B
50
Feed B Condition #2
Feed B Condition #3
40
Feed B Condition #4
30
20
10
0
Total aromatics
Mono
Di
Tri+
Figure 8 Aromatics content in feed and products in Feed B case
As part of the comparison between the 1st and 2nd generation HyBRIM™ catalysts at the same
reactor conditions, we also obtained volume swell data for the two catalysts TK-609 HyBRIM™ and
TK-611 HyBRIM™, enabling a direct comparison of the hydrogenation and hydrodearomatization
difference between them. In Figure 9, the difference in volume swell based on the density reduction
is depicted. Despite a quite low feed aromatic content, TK-611 HyBRIM™ delivers 20% higher
volume swell compared to TK-609 HyBRIM™ at ULSD conditions. This is a significant difference,
and the improvement is remarkable because TK-609 HyBRIM™ is already a high-volume swell
catalyst.
2,4
2,2
Volume swell, vol%
2,0
1,8
20% higher volume
swell
1,6
1st generation
HyBRIM™
1,4
2nd generation
HyBRIM™
1,2
1,0
TK-6 9 HyBRIM™
TK-6
HyBRIM™
Figure 9 Volume swell comparison of TK-611 HyBRIM™ vs TK-609 HyBRIM™ at ULSD conditions
Figure 10 reveals the main reason for the 20% higher volume swell. At ULSD conditions, the product
nitrogen for both catalysts is low and actually below the detection limit. This means that the “monoaromatics saturation high-way” is fully open with the nitrogen inhibitors removed. At such conditions,
the hydrogenation potential of the catalyst is providing the volume swell, and it is therefore observed
that TK-611 HyBRIM™ is saturating an impressive 18% more mono-aromatics than TK-609
HyBRIM™. In refinery terms, the net advantage of TK-611 HyBRIMTM is 660 additional barrels of
liquid product per day for a 40,000 BPSD unit along with better cycle length for the same feed when
compared with TK-609 HyBRIM™.
40
18% higher mono aromatics saturation
seen with TK-6 HyBRIM™
35
Pressure:
SV
30
Aromatics content, wt%
Feed: 70/30 LG/LCO
SG
0.864
S
0.62 wt%
N
696 wtppm
70 barg / 1,015 psi
1 hr-1
25
Feed
20
TK-6 9 HyBRIM™ product
TK-6
HyBRIM™ product
15
10
5
0
Total aromatics
Mono aromatics
Di aromatics
Tri+ aromatics
Figure 10 Aromatics content before and after hydrotreating with TK-611 HyBRIM™ and TK-609 HyBRIM™
Conclusion
The 2nd generation of the HyBRIM™ technology is now available to meet the growing demands for
cleaner fuels and higher volume swell with the introduction of the TK-611 HyBRIM™ catalyst. The
technology offers refiners with high-pressure ULSD units and hydrocracker/FCC pretreaters a
hydrotreating catalyst solution which is second-to-none. The 25% higher activity of TK-611
HyBRIM™ compared to the 1st generation TK-609 HyBRIM™ catalyst can be translated into a lower
reactor temperature while simultaneously obtaining the same conversion of sulfur and nitrogen. This
unlocks the flexibility between longer catalyst cycles, more throughput, better product qualities or
processing of more severe feedstock. Equally important, it is established that the 2nd generation
HyBRIM™ technology exhibits the same high stability that our clients have come to expect from our
BRIM® and HyBRIM™ catalyst series. It has been demonstrated that the TK-611 HyBRIM™ catalyst
has significantly higher hydrogenation activity than the TK-609 HyBRIM™ catalyst due to the higher
activity for nitrogen removal. This will provide higher volume swell to the refiners at similar
conditions, resulting in substantially increased profitability. In conclusion, the advantages of the new
TK-611 HyBRIM™ can be utilized to improve the overall profitability and economy of all hydrocracker
pretreating and ULSD units in multiple ways. Since Topsoe launched this catalyst in April 2016, more
than 10 catalyst charges have been sold in both ULSD and VGO services. Topsoe is committed to
Indian refiners to offer the best catalyst to fulfil their requirements of meeting 100% BS IV diesel by
Q1 2017 and 100% BS VI diesel by Q1 2020 in the most optimal way while improving their refinery
profitability.
Operating experience with the New Continuous Catalyst Withdrawal
System for Improved RFCC Operations
21st Refinery Technology Meet (RTM), Visakhapatnam, India
Paresh Amin & Kate Hovey
Johnson Matthey Process Technologies
E-mail: [email protected]
Introduction:
The dynamic complexity of the FCCU means there are many areas that can be optimized in order to maximize
profitability and reliability. Many refiners focus on optimizing the heat balance through implementing the latest
hardware innovations, such as improving the stripper efficiency or implementing a catalyst cooler. The fresh
catalyst makeup rate and additive usage can also be optimized to complement the heat balance and yield profile
required for maximizing the profitability of the FCCU. It is well recognized that the fresh catalyst and additives
should be injected into the unit on a continuous basis throughout the day. This avoids spikes and dips in the
circulating catalyst activity and ensures smooth operation in order to maximize the unit profitability. In addition,
there is also an economic benefit to be realized from the continuous withdrawal of equilibrium catalyst (E-Cat)
from the unit. The Marathon Petroleum Company (MPC) refinery in Garyville, Louisiana, has led the way in carrying
out process optimizations as a result of continuous catalyst withdrawal, and this article summarizes their
experiences.
The two key considerations in catalyst withdrawals are:
1. Operator safety, withdrawal piping integrity and withdrawn catalyst cooling prior to removal from the
refinery.
2. Stability of operation of the FCCU to allow for optimized heat balance and maximum profitability to be
achieved.
Details of how these are affected by withdrawal practices have been described and represented through
commercial experiences.
The Historical Approach – Batch Withdrawal:
As there is a continuous addition of catalyst into the FCC unit, the circulating inventory is gradually building, which
causes the Regenerator level to increase. On many FCCUs, the Reactor level is normally controlled by the spent
catalyst slide valve and kept at a continuous level to ensure sufficient stripping efficiency. There is always a
quantity of catalyst that is lost from the Reactor and the Regenerator side because the cyclones are not able to
achieve 100% separation efficiency of the catalyst and the product gases or the flue gases. The level of catalyst
losses varies from unit to unit, and the losses from some units can be so significant that there is no buildup of
circulating inventory at all, as the losses are equal to the fresh catalyst make up rates. In an ideal situation however,
the losses are minimal, and the buildup of catalyst caused by the daily additions requires some of this catalyst to
be withdrawn from the Regenerator to maintain the inventory within acceptable limits.
The most common and historical method for catalyst withdrawal is a batch wise procedure carried out once the
Regenerator has reached a specific level. It is a labor intensive exercise that involves the operator opening up the
routing from the Regenerator to the spent catalyst hopper. The hot catalyst (often >700 °C) can be very abrasive,
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especially at high transport velocities, and holing through of the catalyst withdrawal piping is very common and
understandably a significant safety concern.
FCCU licensors will often provide their standard design for catalyst withdrawal piping, which will specify the pipe
class, a section of finned piping for cooling, as well as temperature indications for visibility. It is also common to
see some method of controlling the withdrawal flow without the requirement to partially close the isolation valves.
These include sacrificial orifice plates, or venturis, which will ultimately erode over time and are expected to be
replaced each turnaround cycle. There are some refiners that actually use a number of manual valves in series
which are choked back to control the flow. In this case, once one valve has been significantly eroded the refiner
will move on to the next one, and will continue to do this until the turnaround cycle when all the valves will need
replacement. This is an expensive and unnecessary operating practice.
In most cases, the withdrawal rate is not well controlled and the velocities are both unmonitored and often
u k o due to the e essi e a ie o ooli g ai , which compromises the integrity of the withdrawal piping.
It is common for holes to form especially at elbows and areas of higher velocity. Experience shows that the erosion
is significantly minimized when the withdrawal velocity is kept below 3 meters/second, and the cooling of the
catalyst through the finned sections of piping is greater when the velocities are reduced.
Aside from the above mentioned safety concerns associated with batch withdrawals, the unit stability is also
compromised by the periodic changes in the Regenerator bed level. When the Regenerator bed level is reduced
as a step change, it has an impact on the heat balance, which directly impacts the catalyst circulation and
ultimately the unit conversion. An example of this phenomenon is shown at a US refinery that withdraws
approximately 5% of its unit inventory periodically. The quantity of catalyst withdrawn amounts to approximately
6.8 MT of equilibrium catalyst and the withdrawal takes about 8 minutes. This withdrawal rate is only 3.5% of the
catalyst circulation rate; however, the impact is still significant. The Regenerator temperature increases by 5 °C as
shown in the chart below. The withdrawal period is shown in Chart 1, and the impact on the Regenerator
temperature is shown in Chart 2.
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Chart 1 – US Refiner 1: Batch Catalyst Withdrawal (Bed Levels) 1
Chart 2 – US Refiner 1: Batch Catalyst Withdrawal (Regenerator Temperature) 1
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In addition to the increase in Regenerator temperature, the Regenerator pressure also spikes up during the batch
withdrawal procedure. Charts 3 and 4 show how the Regenerator and Reactor pressures respond during the batch
withdrawal.
Chart 3 – US Refiner 1: Batch Catalyst Withdrawal (Regenerator Pressure) 1
Chart 4 – US Refiner 1: Batch Catalyst Withdrawal (Reactor Pressure) 1
The importance of these changes can be realized by looking at how the product yields have been impacted.
Although the batch withdrawal only results in a temporary period of instability, the economics associated with
these periods are highly significant. The Regenerator temperature and pressure may be able to recover shortly
after the episode, but the same is not realized for the product yields and unit conversion. Chart 5 shows how the
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FCCU slurry yield elevates by 1 wt% during the batch withdrawal and takes twice as long as the withdrawal period
to recuperate.
Chart 5 – US Refiner 1: Batch Catalyst Withdrawal (Impact on Slurry Yield) 1
The displayed impact on Regenerator conditions does not only have an impact on the unit conversion, but also on
the Regenerator combustion kinetics. Excess bed levels will result in longer residence times in the dense phase
and may result in poor air distribution. Regenerator cross sectional mixing can be affected, resulting in coke
combustion issues and localized areas of high temperatures. The same can be said for the low Regenerator levels
following a withdrawal where the dense phase residence time is reduced. Chart 6 below shows an example of
another US refiner who experienced elevated levels of carbon monoxide as the Regenerator level increased. The
carbon monoxide concentration in the flue gas trends closely with the changes in bed level, which additionally
impacted the dilute phase temperatures. Fluctuations like these are not desirable and should be avoided where
possible.
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Chart 6 – US Refiner 2: Batch Catalyst Withdrawal (Impact on Regenerator Emissions) 1
A Novel Approach – Continuous Withdrawal:
Addressing the above mentioned problems associated with batch catalyst withdrawals is made possible by
carrying out continuous catalyst withdrawals. This is a relatively new concept and has been pioneered by MPC in
Garyville, Louisiana, where a continuous Catalyst Withdrawal System (CWS) was installed in March 2016, and is
currently in operation. This system was directly tied into the existing withdrawal piping and is comprised of an
Everlasting isolation valve, a positive displacement fan and three finned pipe-in-pipe heat exchangers to cool the
catalyst, and a collection vessel to receive the cooled catalyst.
The collection vessel uses a sophisticated control logic which is able to carefully control the withdrawal velocities
in continuous operation by using pressure balance between the Regenerator and the collection vessel. The
collection vessel is mounted on load cells so the exact quantity of catalyst withdrawn is known. This results in the
FCCU catalyst balance closure being significantly more accurate, and catalyst loss troubleshooting is made far
easier. The cooled and collected catalyst can be transferred to the equilibrium catalyst (E-Cat) storage hopper
prior to removal from site.
An additional feature of the continuous Catalyst Withdrawal System is that an E-Cat sample point can be included.
The relocation of the E-Cat sample point from the regenerated catalyst standpipe to the Catalyst Withdrawal skid
means that the operator will no longer need to be regularly exposed to high temperature catalyst. Conventional
E-Cat sample points are unreliable and often experience plugging issues, which are a safety concern and require
the operator to manually correct. By allowing the E-Cat samples to be collected after being cooled, the safety risks
are eliminated. An overview of the installation at Garyville is shown in Image 1.
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Image 1 – Continuous Catalyst Withdrawal (Mark-I) Installation – Marathon, Garyville1
The entire Catalyst Withdrawal System can be carefully monitored and controlled through the refinery Distributed
Control System (DCS), giving the operation maximum visibility. A screenshot of the DCS graphic is shown in Image
2 whereby the operator can change the withdrawal set point, see the condition of the equipment, and monitor
withdrawal rates, temperatures and velocities.
Image 2 – Continuous Catalyst Withdrawal (Mark-I) DCS Graphic – Factory Testing Screenshot1
MPC s Garyville refinery has a Flexicracker design FCCU which utilizes an overflow well to define the Regenerator
level. In this design, the batch catalyst withdrawals directly impact the Reactor level, which results in changes to
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the Reactor stripper residence time and the unit conversion. In addition, the insufficient cooling of E-Cat during
the historical withdrawal procedure has limited the catalyst removal truck loading schedule as the hot catalyst
can potentially damage the trucks. The safety concerns associated with the batch withdrawals resulting from
piping erosion have also been experienced in the past.
As a first installation of the continuous Catalyst Withdrawal System (CWS), the cooling capacity of the equipment
far exceeded the original design, and the temperature of the catalyst was decreased to as low as 40 °C at the
lower withdrawal rates. The positive displacement fan operates at maximum capacity so the cooling rate is directly
impacted by the withdrawal rate. Chart 7 shows the collected data at multiple withdrawal rates, ranging from 3
tons/day up to 18 tons/day.
Chart 7 – CWS Cooling Capacity1
During the test phase of the continuous Catalyst Withdrawal System, multiple withdrawal rates were achieved
with a smooth transition between each rate. This showed good control of the system which is dictated by the
pressure balance between the collection vessel and the Regenerator. Chart 8 gives an overview of the varying
withdrawal rates during the test period where the withdrawal rate ranged between 1.5 tons/day and 20 tons/day.
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Chart 8 – CWS Test Period (Withdrawal Rate Variation) 1
Data was collected over a 6-month period where the withdrawal rates were manipulated to control the Reactor
level in this Flexicracker unit. Previously, the Reactor level was difficult to control with the batch withdrawals and
would fluctuate up to 9%; however, after implementing the continuous withdrawal system, a much tighter control
could be achieved. The variation in the Reactor level was decreased from 9% to 5% or a decrease from 4.5%
deviation to 2.5% deviation from the target level. This variation can be further decreased by installing a Reactor
level controller on the DCS that is cascaded to the catalyst withdrawal rate, rather than manually adjusting the
withdrawal rate. Chart 9 shows the continuous catalyst withdrawals over a period of 6 months, with the exception
of a few short outages. Chart 10 shows the significant improvement in the level control with the use of the
continuous withdrawal equipment.
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Chart 9 – CWS Operation over a 6 Month Period1
Chart 10 – Impact on Reactor Level with and without the CWS1
Any variations to the Reactor level will have an impact on the FCCU heat balance, which can significantly impact
the product yields. MPC recognized this, and, with tighter control of the Reactor level provided by the Catalyst
Withdrawal System (CWS), was able to carry out enhanced optimization tests on its Flexicracker FCCU. The Reactor
level was controlled at a range of different set points and the full yield profile was assessed. This identified the
optimum operational Reactor level for this unit and the refinery was able to maintain the Reactor level at this
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point. The economic advantages of being able to do this make the implementation of the continuous Catalyst
Withdrawal System highly valuable and the payback for the project at the MPC refinery was less than 1 year.
Charts 11 and 12 show how the Reactor level impacts the delta coke, represented by the Regenerator temperature,
and how this affects the yield profile (in particular, dry gas yield).
Chart 11 – Reactor Level Impact on Regenerator Temperature1
Chart 12 – Reactor Level Impact on Dry Gas Yield1
A Step Ahead – CWS Mark II:
Since the commissioning of the Mark-I Catalyst Withdrawal System (CWS) at Marathon Garyville, developments
have been made to further improve the efficiency and reduce the costs. The Mark-II design incorporates an
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Page 11
induced draft fan in replacement of the air blower, which is positioned above a patented helical finned pipe design.
This draws the cooling medium (air) through the finned piping and up out of the top of the unit to atmosphere.
The helical finned piping is designed in such a way that it eliminates the requirement for expansion joints which
can be prone to failure and costly to replace. It is a simple, yet effective design which has the added benefit of
significantly reducing the unit footprint. The exchanger skid is only 2.5 x 2.5 meters at the base with a height of 6
meters. See Image 3 which shows the Mark-II setup.
A final improvement which has been incorporated into the Mark-II design is an additional collection vessel. These
essels ope ate i lead a d lag , hi h ea s hile o e essel is olle ti g the ECat, the othe a e discharging
it to the ECat storage hoppe . This t ul akes the Catal st Withd a al S ste continuous ith o e ui e e t
to pause the withdrawals whilst discharging from the collection vessel. The main benefit of this is that it eliminates
the cyclic thermal stresses in the hot piping sections that result from stopping and starting the catalyst withdrawal,
thus improving the piping and equipment integrity.
Image 3 – Mark-II Design Continuous Catalyst Withdrawal System
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Final Summary:
This article has summarized the problems associated with the historical approach to catalyst withdrawals from
the FCCU. The conventional method of batch wise catalyst withdrawals not only gives rise to potential safety
concerns associated with the integrity of the withdrawal piping, it additionally upsets the Regenerator stability
and affects the combustion kinetics. This unsteady period of operation when the catalyst is withdrawn as a batch
can upset the heat balance and cause a degradation of yields which results in an economic loss. It also results in a
change in Regenerator flue gas composition as the combustion kinetics are altered.
An original approach has been developed and implemented at the MPC refinery in Garyville, Louisiana, which
allows the catalyst withdrawal to be carried out on a continuous basis. This system carefully controls the
withdrawal rate such that the withdrawal piping is not exposed to temperature fluctuations or high temperature
catalyst at excessive velocities, which is unavoidable in the historical batch withdrawal approach. It is also much
more efficient at cooling the withdrawn catalyst, which means the removal of equilibrium catalyst from the
refinery can be carried out without any concern of the removal trucks seeing high temperature catalyst.
As MPC s Garyville refinery has a Flexicracker FCCU, the Reactor level is controlled by the withdrawal of
equilibrium catalyst from the unit. Fluctuations to the Reactor level have a significant impact on the unit heat
balance and product yields. The implementation of the continuous Catalyst Withdrawal System (CWS) has allowed
MPC to control the Reactor level more consistently and identify the most optimum level to operate. This results
in an economic advantage that has minimized the payback of this project to less than one year. MPC has led the
way when it comes to optimizing the FCCU operation through stable and continuous catalyst withdrawals and
branched out from the industry norm of the outdated and unreliable batch wise withdrawal practices.
The Mark-II design of the Continuous Catalyst Withdrawal System include recent developments aimed at further
improving efficiency and reducing costs. A patented helical finned piping section is cooled using an induced draft
fan rather than a positive displacement blower which significantly reduced the footprint. The helical finned piping
additionally eliminates the requirement for expansion joints which helped to reduce costs and improve reliability.
And finally, the Mark-II design include two collection vessels which can operate in such a way that there is no
requirement to isolate the catalyst withdrawals during normal operation, improving the stability and reliability of
the withdrawal system.
References:
[1] Fisher, R., E a s, M., Ho ey, K., Hedges, K., Larse , N., & Di kel, B.
7, I pro e e ts i FCCU Operation
through Co trolled Catalyst Withdra als at a Maratho Petroleu Refi ery, AFPM AM-17-45
21st Refinery Technology Meet (RTM), Visakhapatnam, India
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BoroCat™ – An Innovative Solution for Resid Fluid Catalytic
Cracking (FCC) Units
Bilge Yilmaz
BASF Corporation, U.S.A
Lynne Tan Xin Lin
BASF South East Asia Pte Ltd, Singapore
Aaron Liew
BASF South East Asia Pte Ltd, Singapore
Abstract
BoroCat is the latest evolution of BASF’s resid oil Fluid Catalytic Cracking (FCC) catalysts for the
refining market. It is the first commercialized catalyst based on BASF’s new Boron-Based Technology
(BBT) platform designed to optimize refiners’ production yields. BoroCat offers an innovative solution
to refiners processing resid feeds, by minimizing the negative impact of contaminant metals,
specifically nickel. Nickel presents a considerable challenge as it significantly increases hydrogen and
coke yields off the FCC unit, which can eventually result in decreased unit profitability. The enhanced
nickel tolerance is achieved through the novel use of boron, which offers greater flexibility and
profitability to refiners. Successful commercial implementations have verified the ability of BoroCat to
provide improved nickel passivation, leading to demonstrably lower coke, lower hydrogen and
improved bottoms conversion when compared to existing technologies. Case studies from successful
implementations at resid FCC units will be presented to demonstrate the commercial benefits of this
new technology.
Introduction
A significant portion of the FCC units around the world are processing feedstock containing residue
(i.e., resid feeds). The tendency to process resid feeds is higher in parts of the world where access to
light crudes is limited. There are currently many units in Asia, including India, that process resid feeds.
Recent changes in global feedstock quality have been observed in BASF benchmarking studies. In
recent years, specifically 2013 onwards, the percentage of total FCC units operating with resid feed
has increased (Fig.1). This is in stark contrast to 2002, when the percentage of FCC units operating
with resid feed was less than 40%. Globally, more units are processing resid feeds with the average
moving from approximately 40% resid feeds in the early 2000s, to approximately 50% today. This is a
significant increase in the number of refineries that have shifted to resid feeds in the past 15 years.
Figure 1. Refiners globally are processing more resid feeds
Feed quality trends are of interests because resid feeds typically contain contaminant metals
including nickel (Ni), vanadium (V), iron (Fe), and others. These contaminant metals catalyse a variety
of unwanted secondary reactions (e.g., dehydrogenation). Examining metal contaminants, specifically
Ni and V, we can see that the global average is increasing for both of these metals (Fig.2). Asian
refineries especially tend to process heavy feedstocks, resulting in higher metal contamination.
Therefore, Asian refineries have ECat Ni and V levels much higher than the global average (Fig.2).
Figure 2. a) ECat Ni and b) ECat V: Global and Asia averages are both increasing
Ni is a highly active dehydrogenation catalyst, and presents a considerable challenge to refiners as it
significantly increases hydrogen and coke yields. Even small amounts of contaminant metals in the
feed can deposit cumulatively on the catalyst and result in high hydrogen and coke yields during FCC
operation, which is a major concern for refiners. The dehydrogenation effects of Ni can be reduced
through conventional Ni passivation techniques including the injection of antimony (Sb) and the
incorporation of specialty alumina into the FCC catalyst. Ni passivation is especially beneficial when
the FCC unit is on a Wet Gas Compressor (WGC) limit. By removing the WGC limit, a substantial
increase in unit profitability is realized by either increasing FCC unit throughput or severity. While
antimony injection can be effective in reducing the dehydrogenation effect of Ni, it can also poison
Carbon monoxide (CO) combustion promoters and can lead to increased NOx emissions. Specialty
aluminas are also commonly used for trapping Ni by keeping it in a higher oxidation state, and thus
reducing overall hydrogen and coke. However, its efficacy is limited by the very low mobility of Ni and
the immobile nature of the alumina trap. Comparing the amount of contaminant metal concentrated on
the surface versus the core gives the Peripheral Deposition Index (PDI) value
values for Ni are typically high, confirming its very low mobility
1, 2
1, 2
. The measured PDI
. Due to its low mobility under FCC
conditions, Ni is typically concentrated on the outer section of the equilibrium catalyst (ECat) and can
only be passivated when it is in close proximity to the alumina. The low mobility of Ni and alumina in
the catalyst particle leads to insufficient passivation of Ni.
Boron Based Technology (BBT)
Given the limitations of existing technologies, BASF developed a new catalyst technology, Boron
Based Technology (BBT), in which the negative impact of contaminant metals is minimized through
the revolutionary boron based metals passivation approach. The enhanced metals tolerance is
achieved through the loading of the boron compound on a special inorganic support that is introduced
into the catalyst. Combining this novel metals passivation functionality with a pore architecture that
minimizes diffusional limitations of heavy feed molecules, this new technology allows for the lowest
hydrogen and coke yields coupled with higher yield of valued gasoline and light olefin products.
The key feature of BBT is the mobility of boron under FCC conditions and its ability to seek out and
passivate Ni, thus inhibiting its activity in dehydrogenation reactions. Advanced cracking evaluation
(ACE) results of a BBT catalyst deactivated following two different experimental procedures (CMDU
and CPS-3) are presented in Table.1, in comparison with the base case of BASF’s Flex-Tec catalyst.
BASF found that BBT reduces H2 by around 25% and coke by around 13%, or better, depending on
the deactivation method used.
Table 1. ACE results at 75 Wt. % conversion; with 3000ppm Ni and 3000ppm V.
CPS-3
CMDU
Flex-Tec
BBT
Hydrogen
Base
-27%
Gasoline + LCO
Base
+0.8%
Coke
Base
-22%
Hydrogen
Base
-27%
Gasoline + LCO
Base
+2%
Coke
Base
-13%
These laboratory scale results were so good that BASF quickly commercialized this technology. The
first commercialized catalyst from the BBT platform is BoroCat, which was officially launched into the
refining market in early 2016. BoroCat is engineered to provide maximum metals passivation and
superior product yields. This is achieved through the use of Boron Based Technology to passivate Ni
and a proprietary pore architecture is to minimize the diffusional limitations of heavy feed molecules.
In multiple commercial trials, BoroCat has reduced H 2 yield, reduced delta coke, increased gasoline
yield and increased bottom upgrading.
3
BoroCat Commercial Trials
One of the early commercial trials of BoroCat was at a North American refinery in 2014. The FCC unit
runs a moderate resid feed with specific gravity of 0.922 – 0.934, 1-2% CCR, and ECat metals of
~2300 ppm Ni and ~3800 ppm V. The main objectives for the unit are to minimize the amount of
hydrogen produced, as well as to maximize conversion and bottoms upgrading. The refinery wanted
to take advantage of BoroCat to reduce the detrimental impact of nickel and improve operating
flexibility and unit profitability. Prior to the BoroCat trial, the unit was using BASF’s Flex-Tec catalyst,
which contains specialty alumina for Ni passivation.
Figure 3. H2/C1 ratio was lowered at constant equivalent Ni
The hydrogen to methane ratio (H2/C1) from ACE, which is typically considered an indicator of the
hydrogen selectivity of the ECat, was plotted against equivalent Ni level to compare the hydrogen
selectivity between the Flex-Tec and BoroCat samples. BoroCat reduced the H2/C1 ratio by
approximately 20% at constant equivalent Ni level (Fig.3).
Figure 4. Coke yield was lowered at constant equivalent Ni
The enhanced nickel passivation of BoroCat also leads to a significant improvement in coke
selectivity. BoroCat led to a lower ECat coke factor at constant equivalent Ni (Fig.4). This
improvement in coke selectivity can be attributed to the boron compound, which can reduce the
detrimental dehydrogenation activity of Ni and thus the contaminant coke yield. The refinery can
further realize additional benefit from the improved coke selectivity by being able to lower the
regenerator temperature of the unit, thus increasing the catalyst-to-oil ratio and the unit liquid yields.
Figure 5. Unit operating H2/C1 ratio was lowered at constant equivalent Ni
The significant impact of BoroCat in reducing hydrogen yields was also observed in the unit operating
data. BoroCat achieved a 25% reduction in H2/C1 ratio compared to Flex-Tec (Fig.5). The reduction in
hydrogen further demonstrated the improved ability to passivate Ni and reduce the dehydrogenation
activity in the unit.
KBC FCC-SIM modelling was utilized to normalize the operating conditions and demonstrate the yield
and profitability improvements delivered by BoroCat. A summary of the results from the BoroCat
commercial trial is presented in Table 2. At constant operating conditions, BoroCat lowered the
operating delta coke through more effective passivation of contaminant metals, which led to a
significant reduction in contaminant coke. The reduction of contaminant coke allowed a higher
catalyst-to-oil ratio in the FCC unit leading to a higher unit conversion. Despite the higher conversion,
LCO yield was maintained for improved LCO selectivity. Driven by the slurry upgrading and
improvements in H2 and coke selectivity, the bottoms yield decreased by 0.3 vol%. Gasoline and LPG
yields increased by 0.5 vol.% and 0.9 vol.% respectively. As a result of the successful trial, the FCC
unit profitability improvement achieved was $0.32/bbl in winter and $0.38/bbl in summer, based on the
seasonal economic factors.
Table 2. BoroCat commercial FCC unit trial summary
Flex-Tec
BoroCat
Delta
Delta coke (wt/wt)
0.71
0.64
-0.07
Conversion, vol.%
84.9
86.0
+1.1
H2, SCM/BBL
2.47
1.51
-0.96
LPG, vol.%
27.8
28.7
+0.9
Gasoline, vol.%
52.5
53.0
+0.5
LCO, vol.%
12.1
12.0
-0.1
Bottoms, vol.%
7.7
7.4
-0.3
One of the latest successful implementation of BoroCat was at a European refinery in 2016. The FCC
unit at this refinery runs a moderate residue feedstock with typical specific gravity of 0.87, 2.5wt%
CCR, and typical ECat metals of ~4,400 ppm Ni and ~2,500 ppm V. The main objectives for the unit
are to maximize bottoms upgrading, and maximize LPG and gasoline yield. Prior to the BoroCat trial,
the unit was using a competitor’s FCC catalyst.
With BoroCat, the H2/C1 ratio decreased by around 15% (Fig.6), and the coke yield decreased by
around 15% (Fig.6). The reduction in gas and coke translated into an increase in LPG and Gasoline
at constant conversion (Fig.7). Additionally, improved catalyst diffusion characteristics resulted in a
significant improvement in bottoms upgrading to LCO (Fig.8). Furthermore, because BoroCat
delivered improved coke selectivity this unlocked the potential for the refinery to increase unit feed
rate by approximately up to 20%, while maintaining a high unit conversion. Overall, BoroCat
exceeded customer’s expectations by delivering 30% significant increase in bottoms upgrading, 10%
increase in gasoline and LPG yields, 10% reduction in unwanted hydrogen, and 6% reduction in coke
yield.
Figure 6. a) H2/C1 ratio and b) Coke yield: lowered at constant conversion
Figure 7. a) ECat data and b) Unit data: BoroCat improved LPG and Gasoline yields
Figure 8. a) ECat data and b) Unit data: BoroCat improved bottoms upgrading
Conclusion
BoroCat, based on BASF’s new Boron-Based Technology platform, offers an innovative solution to
refiners processing resid feeds, by minimizing the negative impact of contaminant metals, specifically
nickel. Combining this novel metals passivation functionality with a pore architecture that minimizes
diffusional limitations of heavy feed molecules, BoroCat allows lowest hydrogen and coke yields
coupled with higher yield of valuable hydrocarbon products. BoroCat was most recently used by a
refinery in Europe and achieved a reduction in gas and coke, higher LPG and Gasoline yields, and a
significant improvement in bottoms upgrading to LCO. Furthermore, because the coke selectivity was
improved with BoroCat this unlocked the potential for the refinery to increase unit feed rate by up to
20%. BoroCat has been proven in multiple commercial trials globally.
References
1. Vincz, C.; Rath, R.; Smith, G.; Yilmaz, B.; McGuire Jr., R.; Dendritic nickel porphyrin for mimicking
the position of contaminant nickel on FCC catalysts, Applied Catalysis: A, 495 (2015) 39-44.
2. Wu, L; Khalil, F.; Smith, G.; Yilmaz, B.; McGuire Jr., R.; Effect of solvent on the impregnation of
contaminant nickel for laboratory deactivation for FCC catalysts, Microporous and Mesoporous
Materials 207 (2015) 195 – 199.
3. Pan S., Shackleford A., McGuire Jr. R., Smith G.M., Yilmaz B., Creative Catalysis, Hydrocarbon
Engineering. November 2015.
GT-BTX PluS®: GENERATING AROMATICS FROM
FCC GASOLINE
Sachin Joshi
GTC Technology US, LLC
Houston, Texas
INTRODUCTION
A fluid catalytic cracking (FCC) unit is one of the major, and often the largest, units in a refinery.
Various types of FCC (FCC, DCC, RFCC, HSFCC) crack heavy oils to lighter-component streams
that usually fit within the gasoline boiling range. Depending on the different types of FCC, the
aromatics content in FCC gasoline can range from 15wt% to 80wt%. GTC Technology has
developed the GT-BTX PluS® technology, which utilizes the well-known GT-BTX® extraction
process and simple hydrogenation to produce petrochemical-grade Benzene, Toluene, and Xylenes
(BTX) from the FCC gasoline. This is extremely beneficial for refineries looking to convert
gasoline into higher-value petrochemicals.
General Process Description
The conventional process of FCC gasoline treatment is shown in Figure 1.1.1. The full-range FCC
gasoline first goes through a selective hydrogenation unit (SHU) where the light mercaptan sulfurs
are converted to heavy disulfides. The gasoline then is sent to a splitter column to separate light
naphtha (LCN) and heavy naphtha (HCN). The LCN is blended directly into the gasoline pool.
The HCN goes through a hydro-desulfurization unit (HDS) to remove the sulfur (di-sulfides and
thiophenes) before being blended into the gasoline pool.
Introducing the patented GT-BTX-PluS technology.
The FCC gasoline treatment process using GT-BTX PluS® technology has an additional cut of
middle naphtha (MCN). Such middle cut MCN includes high-octane number C6-C9 olefins as
well as BTX aromatics. The FCC gasoline treatment process by GT-BTX PluS® is shown in Figure
1.1.3. The cut point for MCN in the GT-BTX PluS® process is shown in Figure 1.1.4. The MCN
from the FCC gasoline splitter column middle cut is sent to the GT-BTX PluS® extraction unit,
where the aromatics and sulfurs are separated as the “extract”, while olefins and other nonaromatics remain as the “raffinate.”
Feed Definition
GT-BTX PluS® technology demonstrates the flexibility to allow various combinations of
aromatics (5 – 95 wt%) and sulfur (0 – few thousands wppm) in the FCC gasoline with high
performance in one single extractive distillation unit. Depending on the type of FCC, the aromatics
content in FCC gasoline can range from 15 wt% to 80 wt%. And depending on the type of crude,
the sulfur content in FCC gasoline can range from 20 wppm to 1,000 wppm. GT-BTX PluS® can
process all ranges of FCC gasoline alone, or a combination of FCC gasoline with any other
aromatics-rich stream. Table 1 shows typical FCC gasoline quality from a high severity FCC unit.
Description
Density
Unit
kg/m3
Value
677.8
Viscosity
cP
Olefins
wt%
55.2
Aromatics
wt%
24.6
Total Sulfur
wppm
570
TABLE 1
FIGURE 1.
0.2
Example of FCC Gasoline Quality from RFCC
Producing Pure BTX from FCC Gasoline by GT-BTX PluS®
The extract from the GT-BTX PluS® unit consisting of only aromatics and sulfur will be sent to a
separate HDS. There the sulfur will be reduced to <1wppm to meet the petrochemical
specifications; the product after HDS then will be petrochemical-grade BTX. It is an effective and
simple way to convert typically lower-value gasoline to high-value petrochemical product. The
higher the severity of FCC (including DCC, RFCC, and HSFCC), the higher the profit return is
going to be provided by GT-BTX PluS®. Moreover, the olefin-rich raffinate from GT-BTX PluS®
can be recycled back to the FCC unit as an option to increase the propylene yield, converting more
of the gasoline to various petrochemical products. Or the raffinate can be sent to a dedicated recracking process such as Gasolfin.
Simplified Process Flow Diagram of GT-BTX PluS® Unit
GT-BTX PluS® is the application of the well-known GT-BTX® process, but specifically for FCC
gasoline. The process is patent protected in the U.S. and other countries. Figure 1.1.7 shows the
process scheme of GT-BTX PluS® which consists of only two major unit operations: an extractive
distillation column (EDC) and a solvent recovery column (SRC). The EDC cleanly separates the
aromatics, sulfur, and solvent into the bottoms from the solvent-free non-aromatics in the
overhead. The second column strips the aromatics and sulfur in the overhead from the solvent.
The lean solvent is recycled back to the first column through heat exchangers to recover the heat.
A slip stream from the lean solvent is continuously regenerated by a small solvent regenerator to
maintain the solvent quality. The aromatics and sulfur exit the overhead of the SRC column as the
extract product. The olefins and other non-aromatics leave from the overhead of EDC column as
the raffinate product.
FIGURE 2
Simplified Process Flow Diagram for GT-BTX PluS® Unit
Aromatization Process
The aromatization process takes olefinic hydrocarbon streams and produces BTX, with an
aromatic yield approximating the concentration of olefin in the feed. The process technology will
take any olefinic components in the C4-C8 range as feed to produce the aromatics. By-products
are light paraffins and LPG off gases.
FIGURE 3
Simplified process scheme for GT-Aromatization technology
Off-gas
Regen gas
Separator
Product
Separation
Heater
BFW
Reactor A
BFW
Reactor B
LPG and Gasoline
or
BTX
Regen offgas
Feed
The aromatization reaction takes place in a fixed bed reactor; the reactor operates in a cyclic mode
of regeneration (Figure 3). The operation is very simple and it requires no recycle compressor or
hydrogen consumption. Th reactor is operated at 460-540 °C and the pressure is 1-4 bars. The
liquid yield (aromatics) is 47-55% depending on feedstocks. By-products are dry gas and LPG off
gases with yields of 15-20% and 30-35%, respectively. Separation of liquid aromatics products
can be accomplished in the existing BTX recovery post-fractionation unit.
The unit can take the FCC C4 and C5 cuts along with the GT-BTX-PluS C6-C8 raffinate as feed
to add another aromatics increment. This option has the synergistic effect of removing olefins from
the gasoline pool and increasing aromatics production for petrochemical use.
Commercial Experience
GT-BTX PluS® is the application of the GT-BTX® extraction process in FCC gasoline treatment,
to gain special benefits from unique characteristics and requirements for FCC gasoline treatment.
GT-BTX® has more than 50 licensed units as of 2016, and continues to be the most selected
extraction technology for all aromatics-content streams, including FCC gasoline.
GT-BTX PluS® has been applied to a project downstream of a high-severity FCC unit to extract
aromatics and produce petrochemical-grade BTX products. The configuration of this application
is shown in Figure 4. The MSN (where the BTX is) was cut from the FCC gasoline splitter as
middle cut, and sent to the GT-BTX PluS® unit to separate the pure aromatics plus sulfur as extract;
then the sulfur is removed by HDS to <1wppm. The product from HDS is the pure petrochemicalgrade BTX. The LCN and HCN are being treated as regular FCC gasoline and sent to the gasoline
pool. Raffinate, because of its high-olefin content, has an optional route to recycle back to the FCC
unit to produce more propylene. Alternatively, the raffinate could be directly blended to gasoline.
FIGURE 4 Configuration of GT-BTX PluS® Implementation with High-Severity FCC for
producing petrochemical-grade BTX aromatics
The extraction efficiency for aromatics is shown in Table 2
Aromatics Recovery
Aromatics in Raffinate
Aromatics Purity in Extract
Sulfur after HDS
TABLE 2
>99.88%
<0.13wt%
>98.84% “B+T+X”
<1wppm
Performance of GT-BTX PluS® for Aromatics Recovery from FCC Gasoline
Conclusion
For the purposes of producing petrochemical products, it is important to maximize the
petrochemical product yield with the same amount of crude. GT-BTX PluS® is the only technology
that can extract aromatics (BTX) product from FCC gasoline, upgrading it from gasoline value to
petrochemical value. The olefin-rich raffinate from GT-BTX PluS® can also be recycled back to
the FCC or other unit as an option to produce more propylene.
Low Temperature Catalytic Gasification of Petcoke in to
Syngas
Sateesh Daggupati, Sachchit Majhi, Sukumar K Mandal, Asit K Das
Refining R&D, Reliance Industries Limited, Jamnagar
Fossil fuels, primarily coal and natural gas, are the major sources of energy worldwide. I dia’s e ergy
demands are on the rise, annual growth reached 8% recently, and energy demand expected to
nearly double by 2030. This scenario has led researchers to consider alternate sources of producing
liquid fuels, especially from non-conventional sources e.g. petcoke, coal, biomass etc. With the
increasing demand of petroleum and the development of deep resid processing technology of crude
oil thru coking, the output of petroleum coke as a by-product from petroleum refinery has increased
rapidly. Petcoke is solid and it is expensive to transport out of the refinery and It’s a challe gi g task
to utilize the petroleum coke in a reasonable, efficient and clean way since coke is a low value
refinery product. Due to its high calorific value and carbon content compared to coal, petroleum
coke can be used as feed stocks for gasifier to produce syngas. It would be more desirable to convert
petroleum coke into an energy source which is freely transportable in existing infrastructure such as
pipe lines e.g. SNG, syngas or other high calorific value gases. In this context, petcoke gasification has
gained increasing attention in recent years as an alternative source of energy largely throughout the
world. Accordingly, a process for converting low valued petcoke into a more usable gaseous energy
source would be highly desirable.
Although the gasification process is an old known technology, its commercial use has not
been widely exploited throughout the world because of the high costs involved due to
extreme operating conditions and high endothermic heat demand. In order to obtain high
quality synthesis gas, most of the commercial gasifiers (such as entrained flow gasifiers) use
pure oxygen, this demands additional capital and operational expenditures for air
separation units. The process frequently encounters operational problems with reactor
refractory/metallurgy and slag handling issues, etc., because of the severe operating
conditions (T ~ 1400 °C, P > 30 bars). Other commercial gasifiers have been developed based
on the fluidized bed technology in which the carbon conversion is relatively low compared
to the entrained flow gasifiers because of its low operation temperature (i.e. fluidized
gasifier operates in the temperature range between the ash softening and melting point
temperatures). If the gasification temperature in the fluidized bed gasifiers is close to 1000
°C, the ash content of the carbon feedstock starts to soften and the individual particles
begin to agglomerate. The larger sticky particles fall to the bottom of the bed which reduces
the gas permeability and tend to block the reactor and the reactor feed lines and their
removal poses a considerable problem. Generally, both combustion and gasification
reactions occur in the same vessel wherein part of coal/coke gets combusted at the bottom
to supply endothermic heat for gasification that occurs at the upper part of the gasifier.
Several operational issues in a single fluidized bed gasifier are experienced such as
generation of hotspots, agglomeration, etc. If air is used as the combustion agent in the
fluidized bed gasifier, the calorific value of the synthesis gas so produced will be low as N 2
will dilute the synthesis gas.
On the other hand, Gasification of low reactive carbonaceous feedstock such as petroleum
coke is further complicated by its lower gasification kinetics which demands even higher
operating temperature than high reactive carbon feed stock such as lignite and necessitates
the catalytic action for its gasification. Catalytic gasification process has the capability to get
the complete carbon conversion at low temperatures thus by avoiding extreme operating
conditions and issues of caking agglomeration etc. However, in conventional catalytic
gasification, catalyst is impregnated on coal/coke which ends up with major challenging
tasks such as catalyst loss, recovery and regeneration which requires an elaborate and
expensive processing steps.
The afore-said problems can be eliminated by carrying out the combustion and the
gasification reactions in different fluidized bed vessels. The dual fluidized bed process is
capable of producing synthesis gas with air instead of pure oxygen. The novelty and success
of this process scheme depends upon the acceleration of the gasification kinetics of the feed
stock in the presence of the catalytic heat transfer material and the operating conditions
such that the temperature difference between the two zones allows efficient heat transport
by circulation of the catalyst. There is a scope for the development of an efficient catalytic
dual fluidized bed gasification process which can operate at a temperature range where
operational issues such as low carbon conversion, agglomerations, substantial catalyst loss
from the bed, and catalyst regeneration, and the like, are avoided.
A new process has been developed for the catalytic gasification of petcoke in which
combustion and gasification reactions occurs in two different fluidized beds. In the
gasification zone, ~70% of Coke is gasified with steam at a temperature in the range of 700
to 750 deg C in the presence of proprietary catalyst which also acts as heat carrier and
remaining (30%) coke is burnt with air in the combustion zone at a temperature in the range
of 800 to 840 deg C (see figure 1). The exothermic heat is carried to the gasification zone by
the catalytic heat transfer material which is circulating between the two beds.
Figure 1: Low temperature catalytic gasifier
Proof-of-concept studies have been performed on the lab-scale fixed/fluid bed reactor which proved
that the complete carbon conversion can be obtained from the gasification of low reactive petcoke
in presence of a proprietary catalyst at substantially lower temperatures (<750 0C) in minimum
residence time in the range of 2 to 8 min. Several experiments have been carried out to optimize the
process parameters such as catalyst to coke ratio, coke to steam ratio, gasification reaction
temperature, composition of the catalyst (% of active species) etc. In addition, few experiments
were also carried out under similar condition to verify the catalytic activity of other feed stocks such
as sub-bituminous coal, brown coal and biomass etc. The results reveals that the % of alkali in the
catalyst and the ratio of catalyst to feed are the key to get the substantial gasification activity even
for low reactive petcoke at significantly lower temperatures (<750 deg C). The results also proved
that this catalyst has excellent reusability and the molar ratio of H2 to CO in the product gas is ~5:1
(whereas 0.8:1 for commercial gasifier) which crystallizes that this external solid catalyst can also
accelerate the water gas shift reaction.
This catalyst has possessed excellent properties such as hydrothermal stability, attritions resistance
(<7%), high surface area/pore volume, better active metal dispersion and superior gasification
activity. The propriety catalyst remains within the gasifier without losing its activity while feed is
continuously consumed. This is the uniqueness of the process. The proposed process has the
excellent operational features like commercial FCC process (on stream factor ~5 yrs.), and it is
expected to minimize the all operational issues of co
ercial gasifier’s such as catalyst recovery a d
regeneration issues, loss of catalyst along with flue gas etc. The additional advantage of this process
is compatibility with air, i.e. use of pure oxygen is eliminated and the composition of the syngas (i.e.
molar ratio of H2/CO ~5) which leads to the elimination of OPEX a d OPEX for the ASU’s a d CO shift
reactors. This would save minimum 1/3 of the total project cost.
Aromatics Production from Stranded Streams in a Refinery: A
Techno-Commercial Perspective
Hillol Das, Vice President;
Sanjay Rajora, Jt. General Manager;
Manasi Patel, Senior Manager;
Essar Oil Ltd., Mumbai
Introduction
Oil and Gas sector is a very challenging field for its sustainability and enhancement. Refining
business has changed its course over a long period of time due to change in technical,
economic and environmental scenarios across the globe. Refinery configurations are generally
driven by feed specifications, product demand and quality, expansion opportunities and project
budget. With the crudes becoming dirtier and environmental issues enforcing strict fuel quality
requirements, refiners have to adopt robust, deep cleaning and higher conversion technologies
for gaining higher margins. Heavier crudes yield more of heavy oils and bottoms. So, treatment
and utilization of heavier cuts like Vacuum Gas Oil (VGO) and Vacuum Residue (VR) defines
the profitability for a refiner. These days, variety of residue upgradation techniques are
available, which are mature and have proven track records. But even after that there are some
streams in a refinery whose alternative usage can be beneficial and gain higher profits, against
their conventional methods of upgradation.
Stream Identification
We are aware that Fluidized Catalytic Cracker unit (FCCU) is generally part of a traditional
refinery configuration as it mainly contributes in producing Gasoline. It takes sour or sweet VGO
as feed and converts it into products like LPG, Naphtha, Light Cycle Oil (LCO) and Slurry.
Conventionally, LCO is either upgraded to diesel or blended in fuel oil. But characteristically it
has a low cetane no, high nitrogen and aromatic content. Hydrotreatment of LCO is mandatory
for making it suitable for blending in distillate pool. Adding it as a Fuel Oil component further
reduces its value additive capacity. In near future, with the environmental specifications
becoming stringent, getting a place for LCO in fuels may be a costly option for a refiner. For
these reasons LCO is one stranded stream which fetches low value at higher expense.
Similarly, Heavy Naphtha produced from FCC is largely blended in Gasoline. In years to come,
Euro V/VI norms shall pose a problem for streams like FCC Heavy Naphtha, which is an
aromatics rich stream, to get absorbed in gasoline pool. This stream, thus has an alternate path
to the Aromatics complex, making it an opportunity stream in disguise.
So, such potential streams become an attractive feedstock for Aromatics recovery, and the
same is discussed in the following sections.
Project Scheme
An in-house optimization study was taken up to explore utilization of LCO and Heavy Naphtha in
a different manner, and bringing out the maximum value from them. A sophisticated Linear
Programing (LP) tool, historical price sets and in-house financial model have been used for the
evaluation and analysis.
The project scheme targets maximization of Paraxylene (PX) product along with Benzene
recovery. Main feed streams include LCO and Heavy Naphtha from FCC. Hydrogen is
considered as part of feed for hydrocracking/hydrotreating processes. The feed to the complex
is as shown in the below table -
Feed
kTPA
Heavy Naphtha
900
LCO
1300
Hydrogen
30
Aromatics generation follows a proper set of chemical reactions and needs a variety of
equipment and supporting infrastructure, thus making it an Aromatics Complex.
Process path –
LCO is sent as feed to LCO Hydrocracker for maximizing Naphtha. LPG and Ultra Low Sulphur
Diesel (ULSD) products from LCO Hydrocracker are routed to refinery in LPG pool and Diesel
pool respectively.
FCC Heavy Naphtha is fed to Hydrotreater (NHT) for sulphur and nitrogen removal. Naphtha
from LCO Hydrocracker along with treated Naphtha from NHT are fed to a Reformer. Reformate
is further split into Light and Heavy Reformate in a splitter. Light Reformate goes to the Benzene
recovery section for Benzene recovery and raffinate extraction. Thus extracted raffinate goes
back to refinery. Heavy component stream after Benzene recovery goes to Toluene section
where C7s and C8+ are separated. Heavy Reformate and C8+ streams become the feed to
Xylene section. In Xylene section, Paraxylene production is maximized by converting and
isomerizing other C8 molecules. C9/C10+ stream is separated into C9s/C10s and heavier
component. C9s/C10s product undergo Transalkylation with C7s stream to produce C6/C8
components which are again recycled to benzene section for recovering Benzene and
Paraxylene.
The
below
block
flow
diagram
shows
the
Complex
stream
flows
briefly
Raffinate to Refinery
Light
Reformate
Hydrogen
Benzene
Section
BENZENE
Hydrogen
LCO
C6s/C8s
Hydrocracker
H Naphtha
Reformer
Reformate
Splitter
C7s
Toluene
Section
Treated
Naphtha
Heavy
Reformate
FCC Naphtha
Naphtha
Hydrotreater
Trans
Alkylation
C8+
Xylene
Section
PARAXYLENE
C9s/C10s
C9/C10+
Splitter
Aromatics Complex
Heaviers
–
Major products obtained are Paraxylene (PX) and Benzene. Product slate is as shown below –
Product
kTPA
Paraxylene
1058
Benzene
365
Raffinate
539
LPG
91
These yields are generic in nature. More specific yields through technology licensors may result
in further optimization of the product slate.
Price Differentials and Financials
PX-Naphtha and PX-Diesel spread was nearly 400 $/ton in 2016.
Paraxylene Spread with Naphtha and Diesel
1000
in $/ton
800
600
400
200
0
PX CFR SEA
PX-Diesel Spread
PX-Naphtha Spread
Source: Platts & Industry sources
The trend shows that Paraxylene fetches very good margin over Naphtha & Diesel.
A standalone financial model is used to calculate the financials. The estimated capex for the
total Aromatics complex is about 1.1 billon USD. Referring the specific historical prices sets, the
project IRR is calculated to be about 22%. The scheme is found economically viable on various
sensitivities.
Conclusion
Owing to their high Aromatics content, Heavy Naphtha and LCO streams, which otherwise are
routed to Gasoline and Diesel respectively, can have an alternative path for their profitable
utilization. The discussed scheme of producing Aromatics from these streams is seen attractive
on various price sensitivities. Global Market is deficit in Aromatics and Aromatics fetch good
price compared to fuel products. Hence, routing such selected streams for Aromatics generation
opens a new avenue and creates an optimized solution for good value generation. The scheme
can be tested on case to case basis for integration in a refinery.
Disclaimer: The scope of this publication is strictly for knowledge sharing purposes and not necessarily to provide
any recommendation to the audience/readers. Any statement, opinion, or observation made in this paper are those of
the presenter only and do not represent the intent and business plan of Essar Oil Ltd (EOL) nor the figures
represented here are verified by any statutory agency/institution.
Numaligarh Refinery embarks into Bio-Refinery from Bamboo
feedstock -A Game Changer
Author: Geetali Kalita & Rupam Kumar Sarmah
Numaligarh Refinery Ltd, Golaghat, Assam
1.0 INTRODUCTION:
India has already embarked into national mission for alternate fuel replacing the conventional fossil
fuel. The conversion of bio-energy to bio-fuels (Ethanol) and blend with motor spirit to meet the
vehicle emission control strategy is already stipulated by the Government. India contributes only
1% to global fuel grade Ethanol production. Ethanol Blending Programme (EBP) in India envisages
20% ethanol blending and target 10% Ethanol blending in North East.
NRL is flag bearer in support to this national cause and is implementing Bio-Refinery project,
which is a green technology, holding potential to re-juvenile economic growth of the region apart
from fulfilling Government’s mandate for use of alternate fuel.
The NRL Bio-refinery will be based on 3rd generation technology which will enable selective
fractionation of biomass and co-production of multiple products in a sustainable way. To promote
generation of bio-fuels solely from non-food feed stocks, NRL has conceptualized to process
bamboo feedstock to take the first movers advantage of converting Ligno-cellulosic biomass into
Ethanol. Bamboo qualifies as a cellulosic renewable fuel as it matures within 3 to 5 years. NRL will
source bamboo from different NE states. The project will be a Game Changer for the region
The paper highlights project drivers, technology advancement, details of the sourcing &
aggregation model of feed stock, economic model, challenges and environmental & socio-economic
benefits, thus aligning with the national mission of the government on climate change.
1
2.0 PROJECT DRIVER:
2.1 Ethanol Blending Programme (EBP) in India:
In 2002-03, Government of India (GOI) has incorporated 5 % ethanol blending with petrol across 9
States and 5 Union Territories but was only partially implemented due to low sugarcane production.
In September 2008, Union Cabinet approved the National Bio-fuel Policy, making 5% blending
mandatory across all states in the country. However, Government deferred the plan again due to
short supply of sugarcane and sugar molasses in 2008-09.
In 2009, Ministry of New and Renewable Energy(MNRE), GOI adopted the National Policy on
Bio-fuels setting an indicative target of 20% blending of bio-fuels, in HSD and in MS, by 2017.
In 2014, 5 states are blending up to 10% Ethanol (E10). Bio diesel blending started in Aug 2015.
5% ethanol (E5) (from Sugar Molasses) blending in Petrol is implemented in 22 States. The
government’s Major thrust is being given to development of second generation Bio-fuel (non-food
based bio-mass to convert to Ethanol), such as, cellulosic ethanol and algae biodiesel.
2.2 Ethanol production & Usage in India:
In India, Ethanol is mainly produced from sugarcane molasses by fermentation process. 1 ton of
sugarcane gives 70 litres of ethanol whereas 1 ton of molasses yield 220 litres of ethanol. The
Ethanol primarily produced from sugarcane molasses in India caters to the demand of all sectors.
2
As shown in the figure above, the total installed capacity for producing ethanol stands at around 4.7
million tons. Of these the molasses based ethanol production capacity accounts for a dominant
share of 66% in the total whereas the grain based capacity accounts for the rest. The grain based
distilleries producing ethanol are located majorly in states of Maharashtra, Andhra Pradesh,
Madhya Pradesh, Punjab, Haryana and Rajasthan.
There are three main uses of Ethanol in India – potable liquor manufacturing (45%), industrial
alcohol in alcohol-based chemical manufacturing (41%), fuel grade Ethanol for blending with petrol
and other purposes such as use of ethanol as a feedstock to make Ethers (14%).
The demand for Ethanol is estimated and projected for the following grades of Ethanol viz. fuel
grade ethanol, industrial grade ethanol and ethanol going into potable sector. As fuel grade Ethanol
demand is derived from gasoline demand, the first step in estimating fuel grade ethanol is to
estimate and project gasoline (Motor Spirit) demand till year 2030.
3
Demand for E10 ethanol in India was estimated to be ~1.8 million tonnes in FY 14, expected to
grow to ~2.6 million tonnes by FY 20. Base case demand for E20 ethanol in India is estimated to be
~4.3 million tonnes by 2030 with the top 6 states accounting for 51% of demand. However, through
the current molasses route there is a big deficit of ethanol, which requires production of ethanol
through ligno-cellulosic non-food biomass.
Parameter
Present
quantity
Quantity by 2022
(billion litre)
(billion litres)
MS Consumption
29.00
44.00
Ethanol for 10% blending
2.90
4.40
Ethanol available through Molasses route
1.30
1.80
Ethanol Deficit
1.60
2.60
3.0
TECHNOLOGY ADVANCEMENT:
3.1
2ND GENERATION (2G) ETHANOL TECHNOLOGIES:
First-generation (1G) bio-fuels (produced primarily from food crops such as grains, sugar beet and
oil seeds) are limited in their ability to achieve targets for oil-product substitution, climate change
mitigation, and economic growth. Their sustainable production is always under scanner, as is the
possibility of creating undue competition for land and water used for food and fiber production and
possibility of shortage of food. Corn, wheat, and sugar beet can also require high agricultural inputs
in the form of fertilizers, which limit the greenhouse gas reductions that can be achieved.
The cumulative impacts of these concerns have increased the interest in developing bio-fuels
produced from non-food biomass (ligno-cellulosic material), called Second generation (2G) biofuels, which can help solve these problems and can supply a larger proportion of global fuel supply
sustainably, affordably, and with greater environmental benefits.Feed stocks from ligno-cellulosic
materials include cereal straw, bagasse, forest residues, vegetative grasses and short rotation forests.
4
These second-generation (2G) bio-fuels could avoid many of the concerns facing first-generation
bio-fuels and potentially offer greater cost reduction in the longer term.
India generates nearly a Billion tonne of agriculture residues every year, burns nearly 200-300
million tonnes of surplus bio mass – an environment hazard and criminal waste of National
resource. Technologies are now well developed for scale-up. Major challenge is to create a viable
and robust bio mass supply chain. Bio mass aggregation, densification, processing, storage
infrastructure needs to be put in place.
India also wants to use it as a tool to bring much needed stability into agriculture and rural
development. Indian approach is solely non-food feedstock to be raised on degraded, waste and
marginal lands and ensure food and energy security.
3.2
2G Ethanol Technology development:
3.2.1 Dilute Acid Hydrolysis processes:The process was adopted in US during early years for production of ethanol from lingo-cellulosic
feed stock. However, requirement of severe operating condition resulted in de-gradation of glucose
to tars and other un-desirable co-products. Low yield of sugar and subsequent low yield of ethanol
coupled with production of un-desirable co-products means the technology not preferred due to
non-economic production.
3.2.2 Enzymatic Hydrolysis process:
Enzymes catalyze the break-down of cellulose into glucose for fermentation into ethanol. Because
enzymes are highly specific in the reactions that they catalyzed, as oppose to Dilute Acid
Hydrolysis process, formation of by-products are minimized and waste treatments cost also reduces.
Further, enzymatic reaction takes place at milder condition thus giving high yields. Another
advantage is that enzymes are naturally occurring compounds which are bio-degradable. With
advancement of enzyme based technology over the years, ethanol production cost is reduced and
commercial production is economically viable.
5
All conventional Enzymatic Hydrolysis process involves pre-treatment stage as lingo-cellulosic
biomass is naturally resistant to breakdown to sugar. Pre-treatment is required to open up the
structure so that enzymes access the cellulose at ease and reacts/hydrolyze fast and produces higher
yields. There are number of pre-treatment process which includes acid catalyzed steam explosion,
super-critical extraction and dilute acid pre-treatment.
Enzymatic Hydrolysis Biomass Technologies
3.2.3 Advancement in Enzymatic Hydrolysis process:
Presently, more advanced technologies are also available. This technology avoids the main
problems associated with other technologies developed for non-food raw materials and represents a
true third-generation approach for the production of cellulosic sugars and further ethanol. The
technology employ pre-fractionation of biomass in presence of week organic acid, called solvent
which enables co-production of platforming chemicals, such as acetic acid and furfural, In addition,
combustion of co-produced solid bio-fuel (bio-coal) can generate all the energy needed in biorefinery, with some surplus to be used in other production.
The key aspects of advanced Enzymatic Hydrolysis process are:
6

Capable of very high extraction of cellulose to support commercial viability

Ability to convert hemi-cellulose to valuable by-product namely acetic acid and furfural,
which improves project economics/financials considerably

The process separates lignin from cellulose before hydrolysis, which results in very low
enzyme consumption.

The technology employs recycle co-solvent & enzymes by using solvent and enzyme
recovery process, further reducing solvent and enzyme consumption

Low operating pressure & temperature resulting lower Capital & Operating cost

Integrated project module, starting from cellulosic conversion to ethanol along with
production of valuable by-product and co-generation of steam & power.
Advanced Enzymatic Hydrolysis Biomass Technologies
4 BIO REFINERY IN NUMALIGARH REFINERY
4.1
Project Background:
Numaligarh Refinery Limited (NRL) along with Chempolis Oy, Finland is implementing a Biorefinery and are targeting at larger production of sustainable bio-fuels in India, which would reduce
India’s dependence on imported petroleum. The bio-refineries will be based on Chempolis’ 3rd
7
generation bio-refining technology, which enables selective fractionation of biomass and coproduction of multiple products in a sustainable way.
The project aims at construction of a bio-refinery that produces cellulosic ethanol (Fuel Grade) and
platform chemicals from bamboo. The biorefinery will use Chempolis’ formicobio™ technology.
The primary feedstock is bamboo growing in North-East India. Biorefinery may also use other
locally available biomasses such as cereal straws. The biorefinery will consume annually 300,000
tonnes of cellulosic biomasses. The primary products of the biorefinery are cellulosic ethanol and
platform chemical (acetic acid, furfural alcohol). Biorefinery will also produce combustible residues
(biocoal, lignin, stillages). As per the Basic Concept, the bio-refinery will produce 48,900 MT of
ethanol, 11,100 MT of acetic acid, 18,600 MT of furfural alcohol, 160,000 MT of biocoal (to
combustion) and 30,000 MT of stillages. It will be an integrated project module with captive power
generation from bio-residue/Bio-coal for Sales to the grid or on-site consumers.
The project cost is estimated as Rs 950 Cr.
4.2 Project Benefits:
Contribution to Nation’s energy security: The proposed bio-refinery project will
contribute to country’s energy security. The Government of India mandated blending of 5 per cent
ethanol with petrol in 9 States and 4 Union Territories in the year 2003 and subsequently mandated
5 per cent blending of ethanol with petrol in 20 States and 8 Union Territories in November 2006
on an all-India basis except a few North East states and Jammu & Kashmir. This project is a step in
utilizing renewable and environment-friendly sources of energy like ethanol to supplement fossil
fuels.
Employment Generation: The potential of bamboo as an economic resource capable of
generating employment for the rural poor and the skilled and semi-skilled labour in plantation and
others in various value addition activities has remained largely untapped due to lack of an
appropriate policy and institutional framework, covering plantation with community involvement,
technology up- gradation, product and market development.
Sustainability: The direct purchase of bamboo by the proposed bio refinery from the NE states
in the next 3 years’ time will have tremendous impact on the ecology, economy, poverty alleviation
and employment and import substitution. Bamboo, once planted, the clumps will go on producing
8
culms and shoot for about 20 years. In other words, bamboo plantation will act like banks where
people deposit money on fixed term and enjoy the returns in the form of interest.
Value Addition: The added value of a bamboo plantation can be viewed from at least two
perspectives. First there is the financial aspect of generating monetary returns from the cultivation
and use of land. Secondly, the value added should also be considered in ecological terms. Bamboo
is an effective crop for environmental protection. Bamboo produces a full green canopy within three
to four years after planting. So long as selective harvesting is practiced, the canopy will remain
green. This distinguishes bamboo from other forestry species where harvesting implies a reduction
in forest cover.
4.3
Feed Stock Selection & Availability:
4.3.1 Bamboo-a natural Choice of Bio-refinery:
Ethanol can be produced from ligno-cellulosic bio-mass like rice straw, wheat straw, baggase, corn
stover , bamboo , elephant grass etc. All ligno-cellulosic biomass contains cellulose , hemi-cellulose
and lignin. The cellulose and hemicellulose part can be converted to fuel grade ethanol through a
process of enzymic hydrolysis and fermentation in presence of commercial yeast.
Bamboo is potentially an interesting feedstock for advanced bio-ethanol production due to its
natural abundance, rapid growth, perennial nature, low management requirements and better
reduction of carbon footprint compared to an equivalent area of woody plants. The composition of
bamboo is highly similar to other grasses utilized for bio-fuel purposes (e.g. switch grass,
Miscanthus). Its cell wall is comprised of the polymeric constituent’s cellulose, hemi-cellulose and
lignin.
Bamboos are a group of perennial evergreens belonging to the true grass family and enjoying wide
distribution in India, especially in the north eastern region where it is an important resource with
multiple applications.
Compared to other feedstock, bamboo biomass has a relatively high cellulose and low lignin
content which makes it suitable for bio-ethanol production.
.
9
4.3.2 Availability of Bamboo in India:
In India, there are 125 indigenous and 11 exotic species of bamboos belonging to 23 genera. As per
the FAO report on world forest resources, India is the second richest country of the world after
China in terms of bamboo genetic resources. Of India’s total forest area of 67.7 million hectares,
bamboo (both natural and planted) occupies around 13.96 million hectares. This represents 16.7%
of the total forest area of the country and 3.4% of the total geographical area (329 million hectares)
of India. But despite having the largest area under bamboo in the world comprising more than 100
different species, India contributes to only 4% share of the global market. This is mainly attributed
to the low productivity of around 0.4 ton per hectare which is much lower compared to other
countries such as Japan, China and Malaysia which contribute about 80% to the world’s bamboo
market. More than 50% of the bamboo species occur in Eastern India, viz. Arunachal Pradesh,
Assam, Manipur, Meghalaya, Mizoram, Nagaland, Sikkim, Tripura and West Bengal.
In India, 28% of area and 66% of growing stock of bamboo is in the NER and 20% of area and 12%
of growing stock is in Madhya Pradesh and Chhattisgarh. Manipur is the state possessing maximum
diversity of species with 53 species, followed by Arunachal Pradesh with 50 species.
Bamboo by Area %
Bamboo by Growing Stock
In India, most of the bamboo is in the forest land which is difficult to access and thus poor pre and
post-harvest management practices are required. Even though bamboo is grown in the non-forest
land, intensive management is not practiced. It is estimated that only about 15.4% of the total
Bamboo resources of India lie on private lands; as a result, 84.6% of the resources are unavailable
for utilization in industrial purposes without excessive regulation getting in the way (FAO, 2005).
10
Over 39% of the total area under Bamboo is available in the North East Region, which is also the
leader in availability of dense bamboo brakes, in green sound weight and number of green sound
culms.
4.3.3 Bamboo availability in North-East India
North-eastern region of the country is abundant with rich forest resources. The region, which
constitutes only 7.98% of the geographical area of the country, accounts for nearly one fourth of its
forest cover. The total forest cover in the region is 173,219 km2, which is 66.07% of its
geographical area in comparison to the national forest cover of 21.05%.
Northeast region of India is very rich in bamboo diversity. Approximately 60% of the total bamboo
species reported from India is represented from this region. NER alone shares 66% of India’s
bamboo resources. Arunachal Pradesh has the maximum area under bamboo in NER with 16,083
sq. km. followed by Manipur (9,303 sq.km.), Mizoram (9,245 sq. km.), Assam (7,238 sq. km.),
Nagaland (4,902 sq. km.), Meghalaya (4,793), Tripura (3,246 sq. km.) and Sikkim (1,181 sq. km.).
Total bamboo growing stock in NER is 55.14 million MT, among which Arunachal Pradesh has the
maximum growing stock of 14.43 million metric ton (MMT) followed by Manipur (13.73 MMT),
Assam
(12.22
m
MT),
Meghalaya
(7.49
m
MT)
and
Nagaland
(7.27
m
MT).
11
4.3.3.1
Bamboo procurement model for NRL:
To create sustainable and reliable biomass (bamboo) supply chain and to de-risk supply logistics,
NRL will adopt multiple models for sourcing of bamboo, like sourcing from Government agencies,
Contractor/Tendering route and through plantation/cooperative farming.
4.4 Bio-Refinery Technology at NRL:
The Bio-refinery technology at NRL provided by M/s Chempolis Oy, Finland is most suitable for
adoption with bamboo as feed stock as it has also been specially developed for non-food raw
materials (e.g. bamboo, bagasse, straws, oil palm biomass, and other agricultural residues), and it is
based on selective fractionation of biomass with fully recoverable bio-solvent. The key features of
this technology are that it is energy self-sufficient, minimal effluent-generation, low-carbon biorefining technology.
The main components of the process are:
Selective fractionation of biomass- Selective fractionation of biomass is carried out with a
fully recoverable bio-solvent. Fractionation takes place in a much lower temperature and pressure.
12
During fractionation, hemicelluloses and lignin dissolve while cellulose remains insoluble. This is a
key aspect of the technology as it separates lignin from cellulose before hydrolysis. High purity
cellulose requires less enzyme and aids easy & fast hydrolysis.
Washing: After fractionation, dissolved solids and bio-solvent are separated from cellulose by
washing. Washing further purifies cellulose for hydrolysis
Enzymatic Hydrolysis: Enzymatic hydrolysis of pure cellulose into glucose
Fermentation: Conventional fermentation by Saccharomyces Cerevisiae yeasts followed by
conventional separation of ethanol.
Generation of Bio-chemical and full recovery of Bio-solvent: The process recovers
acetic acid & furfural as valuable by-products from hemi-cellulose during fractionation and
evaporation. It also utilizes recycle co-solvent & enzymes through Solvent & Enzyme Recovery
process. Lignin is recovered as dried evaporation concentrate (bio-coal). The recovered lignin is an
excellent solid fuel used for co-gen of steam & power, Bio refinery with this technology will be
Energy self sufficient.
Block Diagram of Bio-refinery at NRL
13
4.5 Environmental Impact:
Gaseous Emission:
The bio-Refining process co-produces bio-coal and stillages that can be combusted at an on-site
power plant, which can produce more heat and power than is needed in the bio-refining process.
Since the biomass feedstock (bamboo) is having negligible sulfur compound and the process does
not employs any sulfur inducing chemicals. As such SO2 emission from combustion of bio-coal is
very less.
Liquid Effluent:
The process recovers and re-uses bio-solvent and water from the process efficiently. Liquid effluent
from the process is very minimum.
Solid Waste:
All the residues from the bio-Refinery will be combusted in the boiler/power plant. The final
residue will be boiler ash containing principally only the ash components of bamboo feedstock.
The boiler ash can be utilized as fertilizer or in cement production.
Self-sufficiency in energy, no net emissions of greenhouse gases:
Therefore, production of cellulosic ethanol with this technology can effectively lead to the
reduction of greenhouse emissions. The diagram below compares greenhouse gas emissions of
different transportation fuels.
14
5
CONCLUSION:
The NRL business models and technological strengths for production of bio-fuels shall be first of
its kind. To promote generation of bio-fuels solely from non-food feed stocks, NRL has
conceptualized to process bamboo feedstock to take the first movers advantage of converting
Ligno-cellulosic biomass into Ethanol. NRL will source bamboo from different NE states. Once
implemented, the project will bring-in sustainable investment in the region, where large section of
local population will be employed. The project will support Government initiative of Ethanol
Blending Program and self-sufficiency in energy. The project will stand as a Game Changer for
the region.
******************************************************************************
Brief Bio data of the Author:
Geetali Kalita, Manager (Tech Services), NRL, delivering service in the Energy section. Her keen
interests are in Energy Optimization, benchmarking study of Fuels Refinery and Alternative and
renewable energy sources for energy security. She is a certified Energy Auditor from BEE .She has
an experience of 17 years in Quality control and Technical Services.
Rupam Kr. Sarmah, Sr. Manager (Tech Services), NRL, leads the Process Technology group with a
special interest in configuration study, Project conceptualization & Project Economics. He has over
19 years of professional experience at NRL in the field of Production & Technical services. He is
also member of Project Steering Committee for Bio-Refinery at NRL.
15
Technical Paper for 21st Refinery Technology Meet, Visakhapatnam
Petroleum/Petrochemical Integration at BPCL Kochi Refinery
By
Shri. George Paul GM (Project Tech- Petchem)
&
Shri. S. Ramanathan CM (Project Tech-Petchem)
1. Background :
BPCL embarked upon a major capacity expansion of its Kochi Refinery from 9.5 to 15.5
Million Metric Tonnes Per Annum (MMTPA) with implementation of the Integrated
Refinery Expansion Project (IREP), which is presently in the commissioning phase. The
configuration of IREP was selected to facilitate the foray of BPCL into petrochemicals by
inclusion of a Petrochemical FCC Unit (PFCCU). About 500 TMT (Thousand Metric Tonnes
per annum) of Polymer Grade Propylene (PGP) would be available from PFCCU for
petrochemical end use. There would also be potential for recovery of about 75,000
MT/annum of Ethylene from refinery off gas stream, another petrochemical building block.
2. Identification of Product Portfolio:
The predominant petrochemical derivative of PGP is Poly Propylene (PP). MRPL and HMEL
have commissioned facilities for PP based upon PGP available from PFCCU and IOCL is also
proposing the same at Paradeep refinery.
BPCL had also similarly initially identified PP as the potential Propylene derivative at Kochi,
during the formulation stage of IREP, as 500 TMT PGP available was sufficient for an
economic capacity PP plant. However, the aforesaid and other projects were already taken
up and there was not sufficient market demand for additional PP capacity.
It was hence decided to venture into the sector of niche/specialty Propylene derivatives
which are predominantly being imported and have high value addition. The major niche
Propylene derivatives are Acrylic Acid, Acrylonitrile, Oxo Alcohol and Propylene Oxide.
Phenol was another option though it is more of a Benzene derivative.
It was recognised that some of the factors characteristic of niche petrochemicals were as
follows:
-
Closely guarded process technology
-
Limited number of producers and technology licensors
-
Marketing capabilities
-
Relatively high capital investment for a low manufacturing capacity etc.
It was also noted that Acrylic Esters (Acrylates) have major domestic demand in the Paints
and Coatings sector. Acrylates are produced from Acrylic Acid and the corresponding Oxo
Alcohol. Acrylates also have high market prices. Hence, it was decided that the proposed
product portfolio for niche Propylene derivatives would be constituted of Acrylic Acid, Oxo
Alcohol and Acrylates.
3. Initial Approach - Joint Venture (JV):
The JV approach
as i itially e isaged for BPCL’s foray i to
iche petroche icals
considering entry into a new business area with associated limitations related to sourcing of
technology, operational experience, need for marketing expertise etc. It was also
understood that an Expression of Interest (EoI) route would not be possible as the
partnership was to be strategic in nature. Accordingly, a scouting exercise was undertaken
for a potential partner by interacting with various Companies through a discussion mode.
Global Management Consultancy firm, McKinsey & Co. provided assistance in this respect.
A potential JV Partner was identified and various activities were jointly undertaken including
preparation of Detailed Feasibility Report, EIA Study etc. Drafting of various JV related
agreements were also commenced. Some of the issues brought up by the proposed JV
Partner were requirement for majority equity holding, competitive feedstock pricing,
marketing arrangements etc.
However, the identified JV Partner decided not to proceed with the project citing their own
internal reasons after jointly working for a period of about 2 years (June 2013).
4. Development as BPCL Project:
It was then decided to proceed with the Propylene Derivatives Petrochemical Project
(PDPP) for the production of niche Propylene derivatives viz. Acrylic Acid, Oxo Alcohol and
Acrylates as a BPCL Project through the conventional technology licensing route. The
following technology licensors were selected for the process units of PDPP:
-
Acrylic Acid: Air Liquide Global E&C Solutions, Germany
-
Oxo Alcohol: Johnson Matthey Davy Technologies Ltd., UK
-
Acrylates: Mitsubishi Chemical Corporation, Japan
Consultancy Fluor Daniel India Ltd. provided assistance for the technology licensor selection
activity.
Some of the aspects that arose during the technology licensor selection process were as
follows:
-
Limited number of technology licensors (many of the major producers do not license
technology)
-
Agreements to be signed before holding Kick off Meeting
-
Patents to be in place considering closely guarded nature of technology
-
Limited number of operating Plants in the case of some technologies etc.
The following Product mix was identified for PDPP based upon the utilisation of 250 TMT of
PGP from IREP as feedstock:
-
Ester Grade Acrylic Acid: 47 TMT
-
Oxo Alcohol (Normal Butanol, 2 Ethyl Hexanol, Iso Butanol): 92 TMT
-
Acrylates (Butyl Acrylate, 2 Ethyl Hexyl Acrylate: 190 TMT
The aforesaid capacities were decided considering economic size of process units and
market demand.
The Detailed Feasibility Report (DFR) was prepared by Engineers India Limited (EIL) and the
Market Survey by Nexant Asia Limited.
The estimated Capital Cost of the project is Rs. 4588 Crore. Approval of the BPCL Board of
Directors was obtained in December 2014 and the Environment Clearance was accorded in
May 2015. The project is expected to be commissioned by end 2018.
It can be termed as a Make in India initiative as it is the first world scale integrated Acrylic
Acid/Oxo Alcohol/Acrylates facility for the products which are being pre-dominantly
imported.
5. PDPP: Integration with Refinery
Though there is a high domestic demand projection for the envisaged niche petrochemicals,
it is essential that the Capex and Opex be optimised to the maximum extent possible by
integration with the refinery facilities. Some of the initiatives in this respect are as follows:
-
Location: PDPP is located in the close vicinity of the refinery premises. Additional land
was acquired across the Public road but the site is being linked to the refinery through
an underpass
-
Utilities: It is proposed that utilities like power, Fuel Gas, steam, treated raw water, DM
water etc. would be sourced from the refinery. Separate Cooling Tower, Refrigeration
unit etc. are being set up in PDPP
-
Synthesis Gas (Syngas): Syngas is a raw material required for Oxo Alcohol Plant of PDPP.
This is to be sourced from Build Own Operate (BOO) Plant being set up along-with IREP
for supply of Hydrogen/Nitrogen
-
Loading Facilities: Existing facilities within the refinery are proposed to be utilised for
despatch of petrochemicals
-
Fire Water System: Integration of fire water system is proposed
-
Effluent Treatment: Final treatment of effluents is proposed in the refinery facilities
6. Applications
The major applications of products from PDPP are summarised below:
Acrylic Acid: Water treatment, Bleaching chemicals, Paper printing chemicals, Dyestuffs,
Paints, Detergent co-builder, textile coating
Oxo Alcohols (Normal Butanol, 2 Ethyl Hexanol, Iso Butanol): Plasticisers, Paint & Resin
formulation, Dyes, Herbicides, textile finishing, Solvents
Acrylates (Butyl Acrylate, 2 Ethyl Hexyl Acrylate): Water and resin based paints, pressure
sensitive adhesives, leather finishing, textile coating, paper coating,
7. Market Aspects/Prices
As per a recent review undertaken by a Committee constituted for MoP & NG, it was noted
that with respect to major imported petrochemicals, the imports of Butyl Acrylate during
2015-16 was approx. 180 TMT and is expected to increase to 300 TMT by 2023-24. The
other envisaged petrochemicals also have high demand and growth rate.
The average price differential between Propylene and Acrylates during the last few years
was about US$ 1000/MT
8. Challenges Faced:
Various challenges were faced during the implementation of PDPP as niche/specialty
petrochemicals is an area with closely guarded technology and limited number of
producers/technology licensors.
-
Unfamiliar nature of Specialty/Niche petrochemical project for BPCL and EIL (PMC) First such project in India
-
Process related aspects due to polymerisation tendency of streams
-
Treatment of effluents with high COD/BOD
-
Procurement related:

Long lead time for delivery of critical items like Acrylic Acid Reactor, Refrigeration
System, Incinerator etc.

Delayed/Non response from Licensor recommended/mandatory vendors – Lack of
knowledge on Indian conditions

Limited number of vendors

Specialised nature/metallurgy of equipment
9. Other Proposals
BPCL proposes to further venture into the niche/specialty petrochemicals area and some of
the proposals in this respect are under study like Propylene Oxide/Polyols from surplus
Propylene, Super Absorbent Polymer (SAP) based upon Acrylic Acid, Ethylene based
derivatives like PVC, EPDM Rubber etc.
10. Downstream Investment
There is good potential for setting up of downstream industries based upon the Niche
petrochemicals proposed by BPCL. The Government of Kerala is co-ordinating the efforts in
this respect.
A New Innovation in Ste
team Methane Reforming - High Effi
fficiency, Zero
Steam Export, Lower Fu
Fuel Consumption
Dr. Siddhartha Mukherjee
ivate Limited
Air Liquide Global E&C Solutions India Priv
A24/10 Mohan Cooperative Industrial Estat
ate
Mathura Road, New Delhi 110 044, India
Introduction
Steam methane reforming (SMR
MR) is the technology very widely used to produ
uce hydrogen from
natural gas and light hydrocarb
rbons. The feedstock is converted in a tubular reactor
re
filled with a
catalyst at high temperatures
es in the presence of steam. The reformin
ing reactions are
endothermic, and the energy fo
for the reactions is provided by heat transfer from
fr
the firebox in
which fuel is burned. The hot re
reformed gas and the flue gas released from the
he firebox are used
for pre-heating the feed, the com
ombustion air and the fuel. In addition, high press
ssure steam is also
produced. This steam is utilise
ed for the reforming process itself while the su
urplus is exported.
The amount of export steam ca
an be adjusted to the user’s needs by process optimisation
o
over a
wide range. Figure 1 shows a ty
typical configuration for a conventional hydrogen
en plant with export
steam.
Figure 1: Typical Conventional SMR
R configuration with Steam Export.
The Concept of Zero Steam E
Export
The energetic efficiency of the o
overall SMR process can be defined by the follo
llowing equation:
Ƞ = (energy of product hydrogen + energy of product steam) / (sum of energy of NG + fuel stream + BFW)
It can be shown that the ove
verall theoretical efficiency of the SMR proces
ess increases with
increasing steam export. Howev
ever, improvement of theoretical efficiency often
n has no economic
advantage. Nowadays, in proce
cess plants, energy integration and optimisation of
o heat exchanger
networks have reached such le
levels that there is very little steam demand from
fr
outside. As a
result, export steam coming from the hydrogen plants fetch only a low economic value.
Therefore SMR units become
e more attractive when they provide the high
hest efficiency for
hydrogen production while min
nimising export steam. This could even mean
an reducing export
A New Innovation in Ste
team Methane Reforming - High Effi
fficiency, Zero
Steam Export, Lower Fu
Fuel Consumption
Dr. Siddhartha Mukherjee
ivate Limited
Air Liquide Global E&C Solutions India Priv
A24/10 Mohan Cooperative Industrial Estat
ate
Mathura Road, New Delhi 110 044, India
steam to zero. Furthermore, iff ssteam is available at a low value, the hydrogen
en plant could even
be designed as a net importerr o
of steam.
A number of options exist to
o reduce the export steam in a standard SMR
R unit. One is the
conventional “Zero Steam Expo
port” plant. For instance using a pre-reformer wit
ith preheat can be
introduced, using high steam/c
/carbon ratio, moderate SMR outlet temperature
ure, maximum feed
preheating and combustion air
ir preheating. Figure 2 illustrates such a configu
iguration. It is more
efficient to utilise the high temp
perature heat of the reformed gas for the endoth
thermic natural gas
reforming reactions rather than
n using it to generate steam. This leads to the
e concept of a heat
exchange reformer, which is the
he basis of the SMR-X technology.
Figure 2: Typical conventional SMR cconfiguration with Zero Steam Export.
The SMR-X Technology
ovel process wherein the heat of the hot reform
med gas is used to
The SMR-X technology is a no
provide further heat for the end
ndothermic reforming reactions. The hot reforme
ed gas flows in an
inner tube arrangement count
nter-current to the feed flow through the cata
talyst bed, thereby
providing a portion of reaction
n heat. In reality, it is a combination of steam reforming
re
and heat
exchange reforming in one com
mpact reformer box (Figures 3 and 4).
Approximately 20% of the energ
ergy required for the endothermic reactions in the reformer can be
provided by this internal heat ex
exchange. The lower temperature of the reforme
ed gas leaving the
reactor leads to significantly low
lower steam production in the process gas boiler
ler. In addition, less
energy has to be transferred ffrom the firebox to the reformer tubes, resultin
lting in significantly
lower flue gas flow and conseq
quently lower steam production in the flue gas boiler.
b
Zero export
steam SMR plants can thus be designed with reduced steam production in both
oth boilers.
A New Innovation in Steam Methane Reforming - High Efficiency, Zero
Steam Export, Lower Fuel Consumption
Dr. Siddhartha Mukherjee
Air Liquide Global E&C Solutions India Private Limited
A24/10 Mohan Cooperative Industrial Estate
Mathura Road, New Delhi 110 044, India
Figure 3: Classical SMR reformer tube arrangement (left) compared to SMR-X technology based on internal heat
exchanger reformer tube arrangement (right)
The Demonstration Plant
Air Liquide has set up a demo plant in Europe to execute long term demonstration tests. The
challenge of such technology is to specify the material grade of the inner heat exchanger tubes
since they are operating in process conditions prone to metal dusting (MD) corrosion risk.
Therefore the design of a reliable heat exchange reformer requires the selection of an
appropriate material for the inner tubes. In addition, an accurate model is required to predict the
gas and metal temperatures in order to assess metal dusting (MD) corrosion potential for the
whole range of the plant’s operating conditions.
The Multipurpose SMR demonstration plant comprises the main process steps of a standard
SMR and is connected to a commercial plant (Figure 5). The demo plant consists of a radiant
firebox consisting of a variable number of commercial size reformer tubes. The reformed gas is
cooled in a process gas boiler and routed to the commercial plant. The flue gas released from
the firebox is cooled in a waste heat recovery section and pre-heats the desulphurised feed gas.
The demo plant is connected to the steam system and the utility network of the commercial
plant. Numerous sensors and sampling points enable calculation of the heat and mass balances
and determination of heat transfer for all sections of the plant.
A New Innovation in Ste
team Methane Reforming - High Effi
fficiency, Zero
Steam Export, Lower Fu
Fuel Consumption
Dr. Siddhartha Mukherjee
ivate Limited
Air Liquide Global E&C Solutions India Priv
A24/10 Mohan Cooperative Industrial Estat
ate
Mathura Road, New Delhi 110 044, India
Figure 4: SMR-X Tube and inner Hel
elical Tube Arrangement
The reformer tubes are equippe
ped with helical heat exchanger coils made of dififferent nickel base
alloys. One of the base materia
rials was selected from the group of lower MD corrosion
c
resistant
alloys, while the second belong
ngs to the group of superior corrosion resistantt alloys.
a
In addition,
some of the inner tube arrangem
ements were equipped with a protective diffusion
n coating.
esults
The Test Programme and Res
The test programme has so fa
far covered more than 10000 hours of operatio
tion at commercial
reforming conditions. At regular
ar time intervals, the inner tubes were inspected
d in
i order to identify
the various MD corrosion steps
ps, from incubation to generalised corrosion. The
Th campaign was
executed covering a wide range
ge of operating conditions.
The demo plant operation wa
was stable for the SMR-X configuration and similar
s
to normal
operation of the reformer tube
be arrangement. The process parameter test matrix
m
was easily
executed over the described br
broad range of process parameters, demonstratin
ting the flexibility of
SMR-X technology and the dem
monstration plant.
A New Innovation in Steam Methane Reforming - High Efficiency, Zero
Steam Export, Lower Fuel Consumption
Dr. Siddhartha Mukherjee
Air Liquide Global E&C Solutions India Private Limited
A24/10 Mohan Cooperative Industrial Estate
Mathura Road, New Delhi 110 044, India
The reformer tube simulation comprises models for external heat transfer from the firebox via
the reformer tube to the catalyst bed and the internal heat exchanger tubes, respectively,
combined with detailed reaction kinetics and flow calculation. Detailed CFD simulation was
applied to represent the reactive flow in the packed catalyst bed as well as the firebox including
the burners. The parameters of the heat exchanger models were calibrated to the measured
packed bed heat transfer prevailing in the reformer tube with the data from the demo plant. The
reconciled simulation model represents the plant data with high accuracy.
The results validate SMR-X technology’s performance in the long term and provide reliable data
sets used for model validation.
Figure 5: Block flow diagram of the Multipurpose SMR demonstration plant; PGB – process gas boiler,
WHRS – waste heat recovery section
Advantages of SMR-X Technology
The SMR-X technology has distinct advantages over the conventional SMR :
1. The internal heat exchange reformer tubes are beneficial for a zero export steam plant
layout. The high temperature heat from the reformed gas is utilised directly for the endothermic reactions. Consequently, less heat is available for steam production. In addition, a lower
amount of fuel gas is required, resulting in lower flue gas flow and smaller convection
section equipment such as heat exchangers and fans. These savings over-compensate the
slightly more complex design of the reformer tubes.
A New Innovation in Ste
team Methane Reforming - High Effi
fficiency, Zero
Steam Export, Lower Fu
Fuel Consumption
Dr. Siddhartha Mukherjee
ivate Limited
Air Liquide Global E&C Solutions India Priv
A24/10 Mohan Cooperative Industrial Estat
ate
Mathura Road, New Delhi 110 044, India
2. Figure 6 illustrates a comp
parison of the CAPEX and OPEX of SMR-X vis-a-vis those of
conventional SMR units with
ith zero steam export for large H2 plants. Both the
th parameters are
lower in SMR-X.
3. Compared to the externall h
heat exchange reformer, the SMR-X is more compact,
c
since all
components are within the
e same box. This leads to lower CAPEX, maint
intenance and total
cost of ownership.
Figure 6 : Comparison of CAPEX and
nd OPEX for a SMR-X vis-a-vis a conventional SMR with Zero
Ze Steam Export
In summary, the SMR-X offe
ffers the following reductions on various para
rameters over the
conventional SMR with Zero Ste
team Export :
•
•
•
•
•
•
CAPEX
OPEX
CO2 – Emissions
Number of Reformer Tubes
Plot Size
Hydrogen Price
5 - 6%
4.0 – 4.5 %
4.0 – 4.5 %
approx. 20 %
approx. 20 %
4.0- 4.5 %
Conclusions
A new SMR technology is prese
sented, which applies reformer tubes with interna
rnal heat exchange.
Internal heat exchange contrib
ributes up to approximately 20% to the energy
gy required for the
endothermic steam methane re
reforming reactions. This advantage is used to propose an SMR
A New Innovation in Steam Methane Reforming - High Efficiency, Zero
Steam Export, Lower Fuel Consumption
Dr. Siddhartha Mukherjee
Air Liquide Global E&C Solutions India Private Limited
A24/10 Mohan Cooperative Industrial Estate
Mathura Road, New Delhi 110 044, India
technology with zero export steam. The plant was operated over a wide range of process conditions, showing the flexibility of the technology.
Finally, the SMR-X Technology, developed in close cooperation with Air Liquide R&D is a
competitive breakthrough for zero steam export as explained in the previous sections. It has an
optimized heat recovery system for substantial hydrogen production cost savings and is
environmentally attractive due to reduced CO2 emissions. In addition, it offers a more compact
design for smaller plot space requirements and reduced CAPEX.
Taking all these factors into account, SMR-X provides a significantly reduced total cost of
ownership for the operating company.
SMR-X Technology is the latest innovative product available in Air Liquide Portfolio for
all the Group´s activities and for our clients.
Excerpts from the ASME PVP 2016 technical paper
An Improved Design of Threaded Closures
for Screw-Plug Heat Exchangers
by Haresh Sippy, MD, TEMA India Ltd
Excerpts from the ASME PVP 2017 technical paper
Design of Threaded Closures for High Pressure Screw-Plug Heat Exchangers
designed to ASME Section VIII Div. 2
by Haresh Sippy, MD, TEMA India Ltd
Hi - Hi
Hi - Lo
The Hi-Lo desig ei g o solete, e ha e o e ted all
Hi-Lo to Hi-Hi i the BORL p oje t, ith the li e so
Che o ’s o se t. The easo s fo o soles e e a e:
. It is - % hea ie / ostlie tha the Hi-Hi desig .
The positio of the fi st affle shifts a d the le gth of
the shell a d tu e u dle i eases due to the
p ese e of a gi th fla ge a d the shifti g of the
ozzle.
. It is u h o e e pe si e to ai tai . Besides, shell
e o al takes lo ge tha u dle e o al a d
e ui es spa e at the ea .
. The studs a e ig i size a d diffi ult to ope . Ofte ,
ou d-headed uts a e p o ided to edu e the BCD.
These e o e i possi le to ope . Se i i g is
diffi ult; effi ie
a d life is edu ed.
. Repla i g the tu e u dle is as good as epla i g the
e ti e e ha ge , as the tu e-sheet is i teg ated ith
the ha el a el.
. it is e haza dous, as the gasketed joi t is likel to
leak to the at osphe e.
Tu e-Sheet to Shell joi t leakage i
a
Hi-Hi
o st u tio .
With
o osi e fluid o tu e side. This
Joi t / Gasket is desig fo a d is
also
tested
at
a
ie t
te pe atu es o l .
It fails i se i e e ause the Spi al
ou d gasket,
ushes u de the
load it's ot desig ed fo , the e t a
load
o
the
gasket
to diffe e tial the
et ee
the Lo
ha
a el
el
is
due
al e pa sio
Allo
all
steel
a d
the
stai less steel i te als. We the
adopt a
filled
etal gasket
ith g o es
ith g aphite epo
as Ka p ofile.
k o
Even though the threaded closures comply with the requirements of ASME Code,
there are still problems of jamming of screw threads due to lack of RIGIDITY. We
have a situation where a pressure vessel which apparently meets all Code rules
and satisfactorily passed the Code required hydrostatic test, seems to malfunction
after a short period of operation.
As ASME gasketed joint designs are based on stress considerations, due to critical
requirements of preventing intermixing of shell side and tube side fluids, an
additional study was made to determine if the existing design would meet the
requirement for leak tightness as per EN 1591-1. Leak tightness becomes a
mandatory requirement in such exchangers which are used in oil refineries to
produce clean fuels.
For assessing the leak tightness of the joint, calculations were made as per EN
1591-1 which is also incorporated in EN 13445 Part 3. It was found that the
applied gasket force was in excess of that required for obtaining leak tightness
under the specified design conditions including effects of progressive distortion
due to frequent re-assembly.
To maintain uniform pressure on the gasket it is imperative that applied bolt load
to be transferred through a RIGID internal cylinder.
Deformation of the Threaded end of the Channel
Figure elo sho s the defor atio of the threaded e d i the radial dire tio ,
of the ha el arrel he the e d o er as su je t to the ha el pressure.
Bell-Mouthi g
It a e see
a al sis that the ag itude of e di g st ess a
a ou t to th ee ti es the a e age lo gitudi al st ess i the u de ut
a ea.
This phe o e o of ell- outhi g a e logi all u de stood a d
suppo ted
al ulatio s to sho that the o i ed e di g a d
lo gitudi al st esses o the u de ut su fa e of the th eads p odu es
a edgi g a tio that esults i adial displa e e t ausi g dilatio of
the ha el.
The efo e, this dilatio that is a ause of edu tio i shea a ea has to
e est i ted
usi g highe all thi k ess i the th eaded egio .
This highe all thi k ess of the ha el a el i
e sel shoots up
the ost of the fo gi g.
TEMA I dia has thus adopted its o
te h olog of added steel to
o e o e this phe o e o of ell outhi g.
Thread Cutting in Progress
A view of channel internals inserted
HHPS Vapour / Treat Gas Exchanger
Summary
RIGIDITY
To prevent leakages and consequent intermixing of fluids, and ensure ease of
maintenance and opening of the plug, all we need can be explained using one word
- RIGIDITY. Based on the type of construction, this RIGIDITY has to be maintained
on various components. This is achieved as follows:
1 All Screw Plug Heat Exchangers
• RIGIDITY o the ha el a el is o tai ed by the Added Steel Ring on the
threaded portion of the channel barrel, thus increasing the wall thickness
2 On Hi-Hi Tube Hi Screw Plug Exchangers
• RIGIDITY of the th eaded po tio of the ha
el a el, as
e tio ed a o e
• RIGIDITY of the gasket that may get deformed or crushed and cause of leakage.
The much-needed RIGIDITY can only be provided by a solid metal gasket with
grooves filled with graphite (Kamprofile) to achieve resilience.
3 Only on Hi-Hi Shell Hi Screw Plug Exchangers
• RIGIDITY of the threaded portion of the channel barrel, as mentioned above
• RIGIDITY of the gasket, as
e tio ed a o e
• As the et diffe e tial p essu e a ts o the tu e sheet i a di e tio agai st the
pressure exerted by the bolts, the internal cylinder across the nozzle openings is
likely to suffer deformation. Therefore, RIGIDITY is mandatory on the internal
cylinder.
Hence, it is imperative to make the cylinder RIGID by way of design that there is
almost no deformation observed in the internal cylinder at the nozzle openings
area. This ensures that the bolt load and the additional load due differential
thermal expansion between the Low Alloy steel channel barrel and the stainless
steel internals is transferred onto the gaskets uniformly to give a leak-proof shell to
the tube-sheet joint.
To download the paper
An Improved Design of Threaded Closures for Screw Plug Heat Exchangers, visit
http://proceedings.asmedigitalcollection.asme.org/proceeding.aspx?articleid=2590
185
The copyrights have been assigned to ASME.
The paper will be available
online December 2017 at
http://proceedings.asmedigitalcollection.asme.org
Connect with me on
www.linkedin.com/in/hareshsippy/
Visit us at
www.temaindia.com
HTRI SMARTPM UNLOCKS RESEARCH FINDINGS, CUTS REFINERIES’ ANNUAL
ENERGY AND THROUGHPUT COSTS
Simon J. Pugh
Heat Transfer Research, Inc.
Heat Transfer Research, Inc. P.O. Box 1390, Navasota, TX 77868 USA
[email protected]
a team of dedicated researchers are testing the fouling
propensity of various crude oils supplied by COFTF member
companies. This team operates three fouling rigs, and their
current focus is to determine how best to translate rig data to
the field (Smith et al., 2017). This work seeks to bridge the
knowledge gap between experimental measurements and the
practical prediction of crude and other hydrocarbons fouling
in refinery heat exchanger networks.
Combining the proprietary, industry-standard HTRI
shell-and-tube heat exchanger analysis methods with
simulation and cleaning scheduling techniques developed in
the Department of Chemical Engineering at the University of
Cambridge, SmartPM is the first validated commercial
software program of its type. It implements tried and tested
dynamic fouling models within a powerful thermo-hydraulic
heat exchanger simulation environment. SmartPM software
was first applied on refinery networks in 2011. After a
rigorous testing and validation process in over 35 networks
around the world, it is now being implemented more widely
by HTRI.
ABSTRACT
With the recent addition of the SmartPMTM software and
consulting team to its roster, leading process heat transfer
technology company Heat Transfer Research, Inc. (HTRI) is
helping its customers identify and achieve multimilliondollar savings in annual energy costs and increased
throughput through rigorous modeling of refinery preheat
trains.
SmartPM applies to full-scale refinery applications the
practical results of decades of crude oil fouling research at
HTRI and other leading research institutions. For example,
dynamic, predictive fouling models for crude oils have been
proven in the field and are now implemented by refineries.
SmartPM monitors and reconciles plant data to enable
engineers to make informed economically driven operational
decisions. These include predicting heat exchanger cleaning
schedules that help reduce maintenance, improve energy
efficiency, improve throughput, and decrease operating
expenses. Very rapid construction and implementation of
network models is possible through the interaction of
SmartPM with Xist®, the industry-standard software for
designing, rating, and simulating shell-and-tube heat
exchangers. Real plant performance data can then be
exported from SmartPM’s built-in data historian to Xist for
design enhancement, revamps, etc.
Using an advanced, highly graphical interface, SmartPM
displays historic and predicted exchanger performance
parameters down to the individual shell level, as well as all
network flows, temperatures, and pressures. The methods
allow for accurate performance estimation, even for older,
poorly instrumented trains.
Crude oil fouling predictions in SmartPM are supported
by the consulting team and the experimental facilities at
HTRI’s Research & Technology Center (RTC) in Navasota,
Texas, USA.
Some benefits of SmartPM are illustrated through
several preheat train case studies.
SMARTPM
–
FROM
PERFORMANCE
MONITORING TO PREDICTIVE MAINTENANCE
The main capabilities of SmartPM include the following.
Rapid network model construction. SmartPM is a
highly graphical software program that uses drag-and-drop
network model-building techniques. It interacts directly with
Xist to allow rapid construction using validated exchanger
data. SmartPM can also export exchanger models to Xist that
are populated with actual, reconciled plant operating data,
extracted from the SmartPM data historian, for one or more
selected days of operation. These models can then be used
for further analysis, such as design enhancement.
Advanced data reconciliation. Unlike conventional data
reconciliation methods that combine energy and mass
conservation with approximate exchanger models, SmartPM
combines detailed heat exchanger simulation at the
individual shell level. In SmartPM, monitoring data are
linked to the network model through the graphical attachment
of tags to appropriate streams. These tags store timedependent flow rate, temperature, and, where available,
pressure measurement data. Data reconciliation takes into
account practical issues such as the paucity of temperature
INTRODUCTION
Leading oil companies from around the world are
collaborating through the HTRI Crude Oil Fouling Task
Force (COFTF) to better understand the mechanisms that
lead to crude fouling in refinery heat exchangers. In the RTC,
1
and flow measurements in many networks, the uncertainties
in thermo-physical properties of the streams, and the
uncertainties in the available measurements.
The data reconciliation approach is based on three steps:
1. Generating missing data through simulation using full heat
exchanger modelling.
2. Filtering unreliable data through a “trusted” heat balance.
3. Grouping heat exchanger monitoring data into periods of
processed crude blends.
In addition to the provision of reconciled temperatures,
flows, and pressures for all locations in the network
(including between shells where no measurements were
made), the key outputs are the historical operational
parameters for the exchangers, desalters, prefractionation
columns, and other equipment. SmartPM reports all the timevarying exchanger parameters that are of interest to process
engineers and detailed exchanger designers, more than 50
parameters in all. The data are provided in tabular and
graphical form. Below is an example exchanger overview or
combination plot showing historic, time-varying values of
three selected key exchanger parameters. Using a common
time scale to compare results, this plot type can include any
of the available parameters for historic and predictive data.
HELPING REFINERS ANSWER OPERATIONAL
QUESTIONS AND GUIDE PLANNING
Process unit engineers and planners at refineries every
day face a number of key questions that the methods of
SmartPM help to address.
Can we predict and manage fouling? The experience of
using SmartPM on multiple crude distillation units (CDUs),
vacuum distillation units (VDUs), and other processes
(Ishiyama et al., 2015) has shown that it is possible to predict
fouling based on previous fouling behavior inferred from
high-quality data reconciliation results (described below).
The models apply to typical time-averaged blends that are
likely to be similar to blends that will be processed in the
future. A very important feature of the preferred crude
fouling model built into SmartPM is that fouling propensity
is expressed by a single numerical number, referred to as the
Fouling Propensity Factor (FPF) (Polley et al. 2011). As
SmartPM is used more widely, a database of FPF values can
be created, and these values can be used for predictions such
as cleaning scheduling and revamping (see below). For
refineries that have crude slates that switch frequently (such
as moving from high sulphur to low sulphur blends),
SmartPM allows the FPF values for various blend scenarios
to be assessed separately.
Which heat exchanger should we clean, when, and
how much would we save? While the use of fouling models
in SmartPM can assist engineers to design or revamp heat
exchangers for lower fouling, the most common and effective
fouling management strategy is cleaning heat exchangers.
Using the FPF value, engineers can explore a range of
cleaning options to account for the cost of heating energy
(gas or furnace oil), the individual cost of cleaning for each
shell, the offline time of each shell, and the minimum profit
that should be made from each exchanger cleaning. All
cleaning calculations account for all the interactions between
offline, online, and newly cleaned exchangers in the network.
The type and effectiveness of different cleaning methods,
assessed from historical data, can be included in the
scheduling. For example, chemical cleaning can be effective
for some shells but not for others, depending on the
temperatures in the shell which, in turn, affect the nature of
the fouling deposit itself. Close examination of historical
fouling data can allow evaluation of the effectiveness of
different cleaning methods.
Some cleaning options in SmartPM are now listed, with
illustrations in the case studies below.
“Top 5” Units To Clean Today for Least Energy Use.
This option ranks the benefits of cleaning individual
exchangers, or units such as shells-in-series, based on the
annual energy benefit. The true cost of an exchanger cleaning
must account for all the interactions between heat
exchangers. For example, the product stream outlet
temperature of an exchanger reduces after cleaning because
the exchanger heat duty increases: colder exchangers
connected upstream of this cleaned exchanger subsequently
have lower heat duties. The network duty gain is therefore
less than the duty gain of cleaned exchangers viewed in
isolation, often by a factor of 0.5. Conventional scheduling
methods cannot model this effect, leading to misleading
Figure 1. Combination plot of the variation in (from the top)
crude-side surface shear, heat duty, and fouling resistance for
a selected exchanger shell. The vertical stripe indicates a
historical cleaning event (here, a hydroblast). The star on the
fouling resistance plot identifies the minimum fouling
resistance after this clean.
Built-in data historian. SmartPM stores all measured
and calculated data. All engineers with access to the
SmartPM project files can rapidly retrieve data. Refinery
engineers and technical services engineers can access the full
data sets, thereby connecting engineers across the oil
company for maximum visibility of plant operations and
simple access to data for detailed analysis.
Network simulation. These capabilities, discussed
below, include heat exchanger cleaning scheduling and
revamps/retrofits for reduced fouling.
2
turnaround, there is often a need to minimize exchanger
cleaning as maintenance costs can be huge and maintenance
tasks are minimized as much as possible. SmartPM can
identify which exchangers to ignore at turnaround, perhaps
by scheduling cleaning before or after or by indicating that
the benefit to the network of cleaning some shells is not
justified at any stage.
What effect will running heavier fouling crudes have
on maintenance? The use of FPF factors for various crude
slates can guide on future operation with crude blends that
have heavier fouling propensities. Even in the absence of an
FPF value for a new blend, increasing the FPF value by, say,
a factor of 2 can indicate possible changes to fouling and
pressure drop buildup in particular shells, as well as changes
in cleaning schedules, which may involve more frequent
cleaning or even cleaning of shells that had previously not
appeared in the schedule. Refinery engineers that have used
SmartPM on several refineries over several operating cycles
(which can result from reviews of historical data) have a head
start in the management of a database of typical FPF values.
What is the effect of revamp/retrofit projects? Through
the use of HTRI’s proprietary heat exchanger design
methods, a full history of all performance parameters is
stored for all exchangers in the SmartPM data historian. By
exporting from SmartPM to Xist to create exchanger files
populated with real plant data, engineers can revamp
exchangers or develop designs with lower fouling propensity.
This can involve designing for higher velocities or lower wall
temperatures, perhaps using proprietary exchanger types and
enhancement technologies. Network pressure drop effects
are often important in this calculation. Exporting new designs
or layout changes back to the SmartPM model, the engineer
can assess (for identical inlet flow conditions over a known
time period) how the network would have performed under
those operating conditions and what the energy savings could
have been. This approach can help justify revamp costs and
increase confidence that the new designs will work as
planned.
Can we increase throughput within current pump and
furnace constraints? Through the use of detailed exchanger
simulation, including pump models and dynamic fouling
models, SmartPM can model the increase of pressure drop
with fouling, in addition to the pressure drop across
exchangers and the entire network as throughput changes.
Similarly, furnace models calculate turndown rates when
furnace firing duty limits are reached. A case study of a UK
refinery preheat train illustrates this latter effect, where
increased throughput and furnace limitations were using
SmartPM (Ishiyama et al., 2013).
Do we have the optimum flow split in our network? For
a network that includes parallel flow paths with uneven heat
transfer resistances, the optimum flow splits are often
estimated using approximate models that do not correctly
account for increased flow resistances due to fouling, or by
trial and error. The detailed exchanger models in SmartPM,
coupled with dynamic fouling models, can calculate the
optimum flow split(s) for either minimum network pressure
drop (which equates to maximum throughput) or maximum
furnace inlet temperature (for minimum energy burn). This
flow split changes as exchangers foul, an aspect that
recommendations. A case study illustrating this effect is
provided below.
Scheduling Cleaning Until Next Turnaround. The use
of predictive dynamic fouling models allows a forward view
of cleaning scheduling that can assist greatly in resource
planning. Heat exchangers are usually cleaned to increase the
furnace inlet temperature and hence reduce energy use and
emissions from burned fuel. An initial estimate of the annual
energy savings for CDUs achieved though correct exchanger
cleaning schedules can be approximated from Figure 2. This
figure is based on multiple examples of preheat train models
around the world, on a consistent energy cost basis. Note that
savings are highly specific to individual networks, the
geometry of the exchangers, and the fouling propensity of the
crude blend (or slate).
Typical CDU Annual Energy Savings (MM USD)
3.5
3.0
2.5
2.0
1.5
1.0
0.5
0.0
0
50,000
Selection of CDUs
Energy cost = $23/MW.hr
$6.3/GJ ; $6.7/MMBtu
100,000
150,000
200,000
250,000
Crude thoughput bbl/day
Figure 2. Typical annual energy savings in range of CDUs
through targeted exchanger cleaning ̶ energy cost basis:
USD 23/MW.hr
However, SmartPM also allows cleaning schedules to be
developed to prevent throughput losses, which can occur due
to network pressure drop limits and/or furnace firing limits
being exceeded.
It can be seen from Figure 2 that cleaning savings
(energy basis) for low fouling propensity crude blends are
related to infrequent cleaning. The importance of selecting
the correct exchangers to clean where fouling is low can be
just as important as for high fouling blends. Refineries with
low fouling may tend to take less care in the analysis of
monitoring data and are prone to making cleaning decisions
that are less reliable. One refinery studied using SmartPM
exhibited low fouling but was forced to reduce crude flows
by half for infrequent cleaning because the bypasses required
to isolate key exchangers had not been installed. SmartPM
studies showed that the installation of bypasses allowed
significant benefits through maintaining production in
addition to saving energy.
Case studies are provided below for two cleaning cases:
cleaning for energy saving and cleaning to maintain
throughput when furnace firing limits are achieved.
Which heat exchangers should not be cleaned at
turnaround? This is the corollary to the previous question,
where heat exchangers are selected for cleaning. At
3
SmartPM can indicate (Ishiyama, et al., 2008). A case study
of flow split calculation was recently developed (Ishiyama et
al., 2017).
Year
Number
of
cleans
CASE STUDY 1: TOP 5 EXCHANGERS TO CLEAN
“TODAY” (ENERGY BENEFIT)
One large US refinery has used SmartPM for several
years to guide their cleaning decisions. Results of a recent
calculation of the best exchangers to clean “now” appear in
Table 1 for only the three most economical cleanings.
1
2
8
8
Unit
Duty precleaning,
MW
Duty postcleaning,
MW
Network
duty gain,
MW
3 med.
hot
shells in
series
1 large
hot shell
2 hottest
shells in
series
11.8
45.2
16.5
10.9
2.5
0.29
13.4
17.6
2.2
0.23
Annual
CO2
saving,
tonnes
Annual
cleaning
cost,
MM USD
Annual
net
savings,
MM USD
8800
18,900
.26
.22
0.72
2.1
Table 2. Summary of annual energy benefits (constant
throughput)
Here, the cost benefits are calculated as the difference
between the predicted furnace inlet temperature without
exchanger cleaning and that with cleaning. Consequently, the
Year 2 figure takes into account energy benefits accrued from
the cleaning work done in Year 1, which is treated as an
investment.
Annual net
energy
benefit, MM
USD
0.97
4.0
Average
furnace
duty
increase,
MW
5.4
11.5
CASE STUDY 3: HELPING REFINERY PREVENT
TURNDOWN
This case study is similar to that of Case Study 2 but adds
the cost of lost production. The cost of lost production is
much greater than saved energy; this cleaning schedule takes
both effects into account.
A large preheat train processes around 130,000 bpd (7
million tpa.) of very variable feedstock, with low and high
sulphur crudes in the blend that frequently switch. The API
of the blend typically ranges from 21 to 26. The FPF of the
average blend is 120, and this value has been shown to model
fouling well over several years of operation, despite the large
variability of the crude blend.
Historically and frequently the throughput has to be cut
as the furnace limit is reached. The refinery now uses
SmartPM to generate cleaning schedules that can be used to
minimize turndown. This is achieved by targeting exchangers
for cleaning to maintain a minimum furnace inlet
temperature, thereby preventing a breaching of the furnace
firing limit.
Figure 3 shows snapshots from the User Graph
component of the SmartPM model. Figures 3a and 3b show
historic mass flow rate and temperature at the furnace inlet,
respectively. In Figure 3a, the historical total mass flow rate
of the crude oil is shown to the left of the vertical line while
the lower (red) solid line to the right of the vertical line shows
the flow rate for a two-year simulation assuming constant,
typical inlet conditions.
The (yellow) circle indicates the condition where the
furnace inlet temperature falls to the critical minimum. After
this point, the furnace model in SmartPM dictates a reduction
in the flow rate to maintain the furnace duty at that limiting
value. Also shown as solid upper (green) lines are (Figure 3a)
the predicted mass flow rates and (Figure 3b) the predicted
furnace inlet temperatures with scheduled cleaning in
operation. Note that the mass flow rate is maintained, apart
from short periods where it drops as exchangers are taken
offline for cleaning. These cleans can be seen clearly from
the predicted furnace inlet temperatures in Figure 3b.
Table 1. Top three shells to clean “today”
For the most cost effective clean (for three medium-hot
shells in series), the duty gain for the exchanger considered
in isolation is 33.4 MW (45.2 – 11.8). An approximate
method, such as those in spreadsheets, might provide an
estimate of a similar magnitude, indicating this to be a very
attractive option. However, an approximate method would
not indicate the network duty gain (that is, the gain for this
unit minus the loss in heat duties of connected units). We can
see that the network duty gain for this cleaning is only around
half that of the unit considered in isolation. Similarly, the
annual net energy benefit of 0.97 MM USD would not be
calculated. This benefit includes the individual shell cleaning
cost, the energy lost when the exchangers are offline, and the
subsequent effect on other units of this clean.
CASE STUDY 2: PREDICTIVE CLEANING
SCHEDULE TO NEXT TURNAROUND: ENERGY
BENEFIT
This same refinery seeks to predict an optimum cleaning
schedule for two years’ operation. This information allows
cleaning budgets and maintenance schedules to be set that
can be justified through predicted economic benefits. While
the schedule below is a calculated optimum, SmartPM also
includes alternative scheduling algorithms, including
cleaning within prescribed cleaning budgets for several
periods (for example, $X for the remainder of the current
financial year and $Y for the entire following financial year).
Table 2 shows the optimum two-year cleaning schedule for
this PHT. Also provided is a detailed summary of the dates
for each clean, which is not reported here.
4
predictions. Its users include engineers managing refinery
process units, technical services groups reviewing revamp
and new design opportunities, and R&D groups studying
fouling.
With cleaning
History
REFERENCES
Smith, A.D., Ishiyama, E.M., Lane, M.R., and Harris,
J.S. (2017). Translating crude oil fouling testing rig data to
the field: a road map for future research. Int. Conf. on Heat
Exchanger Fouling and Cleaning, Madrid, Spain.
Ishiyama E.M., Kennedy J., and Pugh S.J. (2015).
Fouling management of thermal cracking units, Int. Conf. on
Heat Exchanger Fouling and Cleaning, Enfield, Dublin.
Polley, G.T., Ishiyama, E.M., and Pugh, S.J. (2011). Use
of fouling rate models in performance analysis and retrofit
proposal of crude refinery preheat trains, Proc. 12th Int.
Conf. on Petroleum Phase Behaviour and Fouling, London,
UK.
Ishiyama, E.M., Pugh, S.J., Kennedy, J., Wilson, D.I.,
Ogden-Quin, A., and Birch, G. (2013). An industrial case
study of retrofitting heat exchangers and revamping preheat
trains subject to fouling. Proc. Int. Conf. on Heat Exchanger
Fouling and Cleaning. Eds. M.R. Malayeri, H. MullerSteinhagen, and A.P. Watkinson. Budapest, Hungary.
Ishiyama, E.M., Falkeman, E.S., Wilson, D.I., and Pugh,
S.J. (2017). Quantifying implications of deposit ageing from
crude refinery preheat train data. Int. Conf. on Heat
Exchanger Fouling and Cleaning, Madrid, Spain.
Ishiyama E.M., Paterson W.R., and Wilson, D.I. (2008).
Thermo-hydraulic channelling in parallel heat exchangers
subject to fouling, Chemical Engineering Science, 63(13),
3400-3410.
Ishiyama, E.M., and Pugh, S.J. (2017). Effect of flow
distribution in parallel heat exchanger networks: use of
thermo-hydraulic channelling model in refinery operation.
Int. Conf. on Heat Exchanger Fouling and Cleaning, Madrid,
Spain.
Prediction
With cleaning
Figure 3. Top (a) – crude mass flow rate and bottom (b)
furnace inlet temperature
The predicted slopes of flow rate and furnace inlet
temperature reduction match well those from historic
measured data shown in Figure 3. These slopes are based on
calculations of the fouling rate in all exchangers in the
network, using the FPF value of 120. The fouling rates in
each shell will differ, as their geometries and temperature
fields all differ. Additionally, the temperature and wall shear
inside each shell also varies, and these complex effects are
all included in the SmartPM model. The predicted fouling
rates, calculated from detailed exchanger geometries, match
the measured fouling rates well.
Table 3 shows the economic benefits of the cleaning
schedule shown above. As is evident, the throughput savings
dominate.
Year
Cleans
Annual CO2
saving,
tonnes
1
2
12
10
14,400
20,800
Annual
cleaning
cost,
MM USD
0.4
0.32
Annual net
savings,
MM USD
5.7
18.5
Table 3. Summary of annual throughput and energy benefits
CONCLUSIONS
HTRI SmartPM is a valuable tool that allows refineries to
undertake very complex calculations in a simple and easyto-use software environment. Using HTRI proprietary heat
exchanger technology, SmartPM is capable of saving
refineries significant sums through management of fouling.
Its use extends from performance monitoring to performance
5
Aerobic Granular Biomass Technology ‘NEREDA®’–
A Sustainable Wastewater Treatment Option
K. Yagna Prasad
Chief Technology Officer,
VA Tech Wabag Ltd., WABAG House,
200 Feet Radial Road, S. Kolathur, Chennai - 600117
E. Mail: [email protected]
Abstract
Aerobic granulation is seen as the future standard for industrial and municipal wastewater treatment
and subsequently research efforts are quickly developing in this field. As an outcome of a concerted
Dutch program, an aerobic granular biomass technology has been scaled-up and implemented for the
treatment of urban and industrial wastewater. This Nereda® technology is considered being the first
aerobic granular sludge technology applied at full-scale. Operating data from the first municipal fullscale plant confirm the projected advantages with regard to treatment performance, energy-efficiency
and cost-effectiveness. The technology, now applied at tank sizes similar to the world’s largest SBRtanks, is considered proven and applicable for even the largest applications. During the presentation
the latest results and lessons learned will be presented.
Keywords:
aerobic granular biomass, sustainability, innovative biotechnology, Nereda, biopolymer harvesting,
extensive biological nutrient removal, energy efficiency
Introduction
One of the most critical aspects of the activated sludge process has always been the separation of
biomass and treated water. Besides the development of physical separation techniques (membrane
bioreactors) the improvement of settling properties of the activated sludge has been an important
research topic. The basic requirement for biomass with good settling properties is a granular structure
based on compact, dense, large particles with a high specific gravity.
Discovered in 1995 and further developed by Mark van Loosdrecht from the Delft University of
Technology (DUT), the process of using aerobic granular biomass for wastewater treatment has been
scaled up and engineered to suit commercial applications by Royal HaskoningDHV, a Dutch E+C
company and has been commercially branded as Nereda® Technology. WABAG signed a License
agreement with RHDHV for applying this Technology in India and Switzerland.
The Nereda® technology has been applied in various industrial and municipal applications and
demonstrated its robustness and stability. The first full-scale industrial applications date back to 2005,
while in parallel the technology was further scaled-up for municipal application. Following the first
demonstration plants in South Africa and Portugal, a full-scale municipal Nereda® was started up in
2011 at the WWTP of Epe (59,000 PE) followed in 2013 by the WWTP of Garmerwolde (140,000 PE).
On both plants significant improvements regarding process stability, effluent quality (e.g. Epe meets
TN <5 mg/L, TP <0.3 mg/L) and energy savings (>30 %), compared to traditional activated sludge
processes, have been demonstrated. Meanwhile a total of 25 Nereda® plants and 8 process improving
units are in operation or under design, with capacities ranging from 15,000 to 950,000 PE.
sustainable solutions. for a better life.
Page. 1
Aerobic Granular Biomass Technology
Aerobic granules were defined at the First Aerobic Granule Workshop 2004, Munich, Germany which
stated “Granules making up aerobic granular activated sludge are to be understood as aggregates of
microbial origin, which do not coagulate under reduced hydrodynamic shear, and which subsequently
settle significantly faster than activated sludge flocs.”
Starting with activated sludge, aerobic granular sludge can be formed by applying specific process
conditions such as selectively wasting slow settling biomass and retaining faster settling sludge (de
Kreuk et al, 2005). Furthermore, favouring slow
growing
bacteria
such
as
Poly-phosphate
Accumulating Organisms (PAOs) has been shown to
enhance granulation (de Kreuk et al, 2006). Aerobic
granular sludge consists of bio-granules, without
carrier material, of sizes typically larger than 0.2 mm.
The granular biomass can be used to biologically
treat wastewater using similar processes to activated
sludge system, however the granular sludge has a
distinct advantage of faster settling velocities when
compared to activated sludge, which allows for higher
reactor biomass concentrations (e.g. 8-15 g/l) (de
Kreuk et al, 2007). SVI5 of aerobic granules being
comparable to SVI30 of activated sludge. Figure 1
illustrates the settling properties of aerobic granular
Figure. 1: Comparison of SVI5 of Activated sludge & Nereda
sludge compared to activated sludge after 5 minutes
of settling. Furthermore the particles formed provide a structured matrix for biomass growth, containing
spheres with anaerobic, aerobic and anoxic conditions (de Kreuk et al, 2007) which are populated by
different microorganisms including nitrifiers, denitrifiers, and glycogen accumulating organisms (GAO)
along with PAOs.
This allows for a simultaneous execution of the processes required for nutrient removal, and provides
the foundation for a process that is both simple and requires minimal space.
Figure 2 shows a pictorial representation of the distribution of biological organisms within aerobic
granules compared to activated sludge. As we see, compared to normal activated sludge there are
much more nitrifiers, PAO’s and
Activated sludge from SBR
Aerobic Granular Biomass
GAO’s. This explains the superior
BNR performance noticed in Aerobic
Granular Reactor (AGR). Also this
picture
illustrates
why
it
is
experienced that granular biomass is
more stable and less sensitive
towards toxicity and fluctuations. The
bacteria/water surface is much
smaller and toxicity will mainly effect
the bacteria at the outer shell. The
higher bacteria population makes
them to recover quickly. More
importantly, especially the GAO and
Figure 2: Difference between Activated Sludge & Granular Sludge
PAO produce Excellular Polymeric
sustainable solutions. for a better life.
Page. 2
Substances (EPS) or biopolymers that actual form the backbone of the granule and are the house for
the microorganisms. This backbone is chemically very stable. By the way, recently research has started
how in future to recover the biopolymer as valuable byproduct.
The structure of aerobic granule is depicted in Figure. 3
Oxygen gradient in granule
Simultaneous COD, P and N-removal
Heterotrophic Organisms
Anaerobic zone:
 Nitrate reduction to nitrogen gas
 Phosphate removal
COD + NOx + PO43-  N2 + CO2 + H2O + poly-P
Aerobic zone:
 Biological oxidation
 Ammonium oxidation to nitrate
COD + O2  CO2 + H2O
NH4 + O2  NOx
Ammonia Oxidizing Organisms
Figure 3: Structure of Aerobic Granular Biomass
When aerated, an oxygen gradient forms within aerobic granules whereby the outer layers are aerobic
and the inner core is anoxic or anaerobic (de Kreuk et al, 2007). Nitrifiers and heterotrophic bacteria
proliferate in the aerobic outer layer of the granules, enabling the degradation of organics (COD
removal) and nitrification (conversion of ammonia to nitrite/nitrate) respectively (de Kreuk et al, 2007).
A simultaneous nitrification-denitrification process occurs whereby the formed nitrates (from
nitrification) are denitrified (conversion of nitrate to nitrogen gas) in the anoxic core of the granules
(Pronk et al, 2015). PAOs in the aerobic granules enable enhanced biological phosphorus removal
whereby phosphate uptake occurs during aeration and phosphate rich waste sludge is subsequently
removed from the system (de Kreuk et al, 2005). Aerobic granular sludge can therefore achieve
biological nutrient removal in a single tank without the need for separate anaerobic and anoxic
compartments or tanks. Comparatively, activated sludge systems capable of biological nitrogen and
phosphorus removal require at least 3 tanks or zones (anaerobic, anoxic and anaerobic) and multiple
recycles between the zones or tanks (Wentzel et al, 2008).
Nereda® systems are preceded by conventional pre-treatment consisting of screening, grit removal
and, depending on the application, FOG (fats, oils and greases) removal; whilst primary sedimentation
is optional. Typical reactor depths range from 5.5 to 9 m, with lower and deeper depths possible; whilst
secondary settling tanks and major sludge recycles are not required for the Nereda® system.
sustainable solutions. for a better life.
Page. 3
The Nereda® Process
The Nereda® process uses an optimized sequencing batch reactor (SBR) cycle in which the 4 steps of
a typical SBR cycle are reduced into 3 steps (Figure 4):
®
a) A Typical SBR Cycle
1. Fill, 2. Aeration, 3. Settling, 4. Draw,
b) Nereda Cycle
1. Simultaneous fill / draw, 2. Aeration, 3. Sedimentation
Figure 4: SBR Cycle & Nereda® cycle
1.
Simultaneous fill/draw: During this stage the wastewater is pumped into the reactor and at the
same time the effluent is drawn.
2.
Aeration: During the aeration phase, biological conversion take place. The outer layer of the
granules are aerobic and it is here where nitrifying bacteria accumulate. This forms nitrate that is
then denitrified in the anoxic core of the granules. In the final step phosphorous uptake occur
3.
Sedimentation: Following the biological processes, a sedimentation phase separates the clear
effluent from the sludge. The time for phase separation is short due to the excellent settling
properties of the sludge. The system is then ready for a new cycle.
The key advantages of Nereda® are summarized as follows:
Cost-effective
Compact and uncomplicated tank design

Less mechanical equipment

No separate clarifiers needed

Easy to operate

Robust and reliable process performance

Fully automated plant operation possible
Sustainable

High effluent purity and efficient nutrient removal

No or minimal use of chemicals

Significantly lower energy consumption
sustainable solutions. for a better life.
Page. 4
To meet the effluent demands and energy
efficiency requirements of the WWTP,
optimisation of the Nereda® process can be
controlled by online process analysers
measuring ammonium, ortho-phosphate,
oxygen, and the oxidation reduction potential
(ORP).
For
less
stringent
effluent
requirements, typically the main process
control parameters are oxygen and ORP. Like
for all advanced controls it is desirable that the
measurement values are highly reliable.
At the Epe wastewater treatment plant reliability
of the ammonium, phosphate and nitrate
measurements is ensured by a predictive
diagnostic system called Prognosys that
monitors and interprets the instrument’s internal
signals to inform the user of the instrument
condition. The reading is expressed as a
percentage value and is designed to inform
operators about upcoming maintenance needs
before measurements become questionable
and might affect the process.
Ammonium and phosphate are measured using
the outdoor versions of the Amtax sc (NH₄+) or
Phosphax sc (PO₄3-) analyser respectively.
These analysers do not measure directly in the
process medium, but the sample for analysis is
pumped from the Nereda® reactor, pre-filtered
(<0.45 microns) in a self-cleaning module and
transported to the analyser. Both analysers
have an analysis time of approximately 5
minutes. Oxygen (LDO sc) and pH/ORP (pHDS sc) sensors can be directly placed in the
medium thereby delivering measurement
values in real time.
In figure 5 the different measurement signals of
sensors and analyzers during the aeration
phase are shown as a trend line. It can be seen
that during the aeration cycle, the oxygen
concentration is kept constant and there is a
decrease of the ammonium and the orthophosphate concentration. The ORP signal
increases in accordance with the increasing
ratio of oxidised to reduced species.
sustainable solutions. for a better life.
At Epe and Garmerwolde, a significant
parameter for process control in the aeration
phase of the Nereda® process, is the NH₄+
concentration value delivered by the Amtax sc
analyser. The reliability of the NH₄+ and other
measurement values is constantly monitored by
Prognosys and classified in percentage values
as the so-called measurement indicator. In case
Concentration, mg/l
Controlling the Nereda® process
0
0.5
1.0
NH4-N
1.5
PO4-P
2.0
O2
2.5
ORP
3.0
Time, h
Figure 5: Trend lines from the online measurements
during the aeration cycle of the Nereda® process
the value of the measurement indicator starts
decreasing from 100% there is still enough time
to take action before results get questionable. If
the value should drop below 50%, an alternative
strategy to control the aeration is activated
using the mV value delivered by the ORP
sensor as backup signal.
Data transfer and communication
All measurement signals from one reactor are
captured by a single SC1000 controller. TCP/IP
is used for the communication between the
controller and the AquaSuite® Nereda® PLC.
Controller and attached instruments can be
remotely monitored via the network, i.e.
measurement values as well as the status of the
instruments provided by Prognosys can be
retrieved and maintenance steps like a
calibration can be remotely started.
Page. 5
Results from Epe WWTP, The Netherlands
Epe WWTP is a full scale Nereda® plant which was designed and constructed by Royal HaskoningDHV
in 2010-2011 and has been operational since September 2011. The plant consists of the following
main processes; inlet works with screens and grit removal, followed by three Nereda® reactors and
effluent polishing via gravity sand filters. The Nereda® reactors are designed to take average daily
flows of 8,000 m3/day and a peak flow of 36,000 m3/d. The waste sludge is thickened via a gravity belt
thickener and transported off-site. The performance of the plant is outlined in table 1.
Table 1: Epe WWTP – Performance Results during Process Verification
Effluent – average [mg/L]
Parameter
Influent [mg/L]
COD
879
27
32
BOD
333
<2.0
<2.0
N-Kjeldahl
77
1.4
1.8
NH4-N
54
0.1
0.1
<4.0
5.1
N-Total
Effluent (95%ile)
P-Total
9.3
0.3
0.34
Suspended Solids
341
<5.0
<6.0
One key advantage of Nereda® is reduced power consumption. At Epe, the original plant energy consumption
was approximately 3,500 kWh/d. With Nereda®, the average daily consumption is now 2,000 kWh – 2,500 kWh.
This is approximately 35% less than all types of similar sized conventional plants in the Netherlands.
Conclusions
Existing Nereda® plants demonstrated that the technology is capable of effectively treating wastewater
for removal of ammonia, total nitrogen and phosphorus. The process is effective at removing these
parameters to low levels, in line with future effluent consent limits that might be put in place by the EU
water framework directive. Notably, the technology is delivering wastewater treatment at a significantly
reduced CAPEX (plant size, footprint) and OPEX (energy, chemicals) compared with conventional
technologies on the market.
References
de Kreuk, M.K., Heijnen, J.J. and van Loosdrecht, M.C.M. (2005) Simultaneous COD, Nitrogen and
Phosphate Removal by Aerobic Granular Sludge. Biotechnology and Bioengineering 90 (6), 761-769.
de Kreuk, M.K., and van Loosdrecht, M.C.M. (2006) Formation of aerobic granules with domestic
sewage. Journal of Environmental Engineering 132 (6), 694-697.
de Kreuk, M.K., Kishida, N. and van Loosdrecht, M.C.M. (2007) Aerobic granular sludge – State of
the art. Water Science Technology 55, 75-81.
Pronk, M., de Kreuk, M.K., de Bruin, B., Kamminga, P., Kleerebezem, R., and van Loosdrecht,
M.C.M. (2015) Full scale performance of the aerobic granular sludge process for sewage treatment.
Wentzel, M.C., Comeau, Y., Ekama, G.A., van Loosdrecht, M.C.M., and Brdjanovic, D. (2008)
Chapter 7: Phosphorus Removal. In: Biological Wastewater Treatment: Principles, Modelling and
Design (ed. Henze, M., van Loosdrecht, M.C.M., Ekama, G.A., and Brdjanovic, D.), IWA, London.
sustainable solutions. for a better life.
Page. 6
Refinery’s water sustainable assessment; The need of hour
Jayant Kumar Joshi, Head Sustainability, Engineers India Limited, [email protected]
SN Sukhwal, Adviser (Technical) Centre for High Technology, [email protected]
Abstract:
Going forward reduction of Carbon foot print and reduced dependence on fresh water is going to
be guiding operational philosophy. With fresh water becoming scarce, due to increase in
population & per capita consumption, refineries would need to resort to alternate routes of water
consumption such as rain water harvesting, recycle and reuse of treated waste water, treat MSW
(Municipal sewage water) primarily & seek sea water to meet its operational requirement.
Water consumption is closely associated with the water footprint assessment of any refinery. This
is an analytical tool, it can be instrumental in helping to understand how activities and products
relate to water consumption and pollution and related impacts and what can be done to make sure
activities and products do not contribute to unsustainable use of freshwater. As a tool, a water
consumption assessment provides insight to a process unit for water utilisation, and help to
understand what can be done to optimize the requirement of water as a whole in a system.
A full water assessment consists of four distinct phases
1. Setting goals and scope.
2. Present status of water consumption in the refinery
3. Water sustainability assessment.
4. Area of concern and action plan.
Analysis of present practices surely help in integrating usages of water in the process, say for
example, use of optimum quality of water at desalters by way of mixing SWS/ DM water etc by way
of water pinch methods.
The water assessment not only includes data collection of source/ quality, flow rate and seasonal
pattern, treatment type/Treatability; designated use, specific consumption and waste water
generation figures from process units based on capacity of the plant, recycle reuse options etc.
This article reviews all latest technologies such as low temperature multi stage distillation, vapour
compression, Electrodialysis, ultrafiltration & reverse Osmosis for the water management &
examines how different techniques may be useful in these areas in meeting the most stringent
standards in reducing the overall water requirement of a refinery. No chemical treatment in cooling
water circuit such as magnetic treatment, pulse power technology and Galvenic cell technology are
gaining importance as they require less bleed off and less fresh water requirement.
For oil removal the new technologies have targeted chemical bonding, reusable micron filtration,
and media adsorption and coalesce as techniques to arrive as a niche product, which can provide
answer to the present & future requirement of waste water treatment.
This paper shall provide some real cases of water assessment in the context of Indian refinery
where various sources such as sea cooling water network, Raw water reservoir, Bearing cooling
water system, condensate reuse, problem in using refinery treated waste water are discussed with
a view to optimize requirement of fresh water in processing units.
Introduction
In view of the limited availability of fresh water resources and the need for their conservation, the
implementation of water recycling concepts within the framework of sustainable water
management strategies is of crucial importance. Realizing the deteriorating water resources and to
give effect to environmental measures and policies for pollution control, various steps have been
initiated by the Government. In India, the regulatory requirements in terms of Environmental
protection are quite stringent. By any benchmark, India has an extensive environmental
management system with a comprehensive set of environmental laws, specific statutory
mandates, regulatory instruments, and institutional frameworks to implement and enforce
environmental policy objectives. The regulatory authority, viz. Central Pollution Control Board
(CPCB), has proposed stringent minimum national standards for oil & gas industry. Besides, the
industries are required to meet the site specific standards which form part of their environmental
clearances.
Despite taking all these measures, Refineries need to relook at their water utilization capabilities.
Refineries utilizes large amount of water in various processes as utility and they also generate
huge amount of waste water. Significant quantities of water—primarily for processing and
cooling— are needed to produce fuel in a refinery. India's current refining capacity is 230 MMTPA,
including the just commissioned 15 mmtpa IOC refinery at Paradip. As a rule of thumb Refineries
use about 2.5 gallons of water for every gallon of crude processed, meaning that India, which
refines nearly 75 billion gallons of petroleum products annually, consumes about 190 billion
gallons of water to produce fuel.
Water management has always remained at the core of process optimization to reduce fresh water
requirement and waste water generation. Availability concerns of fresh water and more stringent
effluent discharge requirement have necessitated increased focus on water management in a
refinery.
Water used to be seen as a low-cost resource to refineries, and was used inefficiently. However,
as the standards and costs for wastewater treatment increase and the costs for raw water makeup
increase, the refineriey has become more aware of water costs. In addition, large amounts of
energy are used to process to move water through the refinery. Hence, water savings will lead to
additional energy savings. This paper attempts to address all such issues.
Water assessment audit
A water audit is an excellent way to understand refinery’s current water use and future water
savings. Generally, a water audit provides a detailed description of water usages, identifies
potential water and financial savings, and recommends various water efficiency upgrades.
Additionally, the world business council for sustainable development WBCSD’s Global Water Tool
is an easy-to-use tool to map refineries water use and assess risks relative to their operations and
supply chains.
The first step in this direction is to find out the total water consumption in a refinery. Since, water
consumption is closely associated with the water footprint assessment of any refinery that can be
instrumental in helping to understand how activities and products relate to water consumption and
pollution and related impacts and what can be done to make sure activities and products do not
contribute to unsustainable use of freshwater. As a tool, a water consumption assessment
provides insight to a process unit for water utilisation, and help to understand what can be done to
optimize the requirement of water as a whole in a system.
A full water assessment consists of four distinct phases
1. Setting goals and scope.
2. Present status of water consumption in the refinery
3. Water sustainability assessment.
4. Area of concern and action plan.
Analysis of present practices surely help in integrating usages of water in the process, say for
example, use of optimum quality of water at desalters by way of mixing SWS/ DM water etc by way
of water pinch methods.
Setting goal and scope
Setting goal and scope is one of the key activities in a full refinery water assessment. It is
important to define the purpose of carrying out water assessment and intended output. The overall
scope should be around following points:

Raw water treatment facilities in the refinery

Quantity & quality of water used in a particular process

Waste water generation from the unit

Internal recycle of waste water

Waste water treatment facilities

Waste water recycle and reuse alternatives

Preparing a water mass balance for the refinery

Formulation of Zero effluent discharge philosophy
The goal & scope are transformed in the form of a questionnaire, where the intended information is
asked sequentially from the end users in terms of questions. It may require efforts from different
groups such as technical, HSE, operations etc to get involved to fill up the required information
from the questionnaire in a designated format. To coordinate among different groups a project
coordinator is designated. It will be the role of project coordinator to understand the requirement
asked in the questionnaire vary clearly and dissipate information to the group.
Present status of water consumption in the refinery
Once the questionnaire has been received by the various groups, the question pertaining to their
area of activity should be reported. It is quite possible that most of the information is readily
available in terms of primary data but in some cases refiners are required to generate information
through analysis and information on past data/ records. Sometimes, the role of operators is of
prime importance because e.g they are the one who are quite aware about the frequency of
draining, their characteristics and flow from a particular drain. In some cases where there is no
access to drain and it is difficult to determine flow the importance of experience of operators are
helpful.
The water intake and treatment facilities shall involve collection of data with respect to Raw water
quality of the respective water sources
 Identification of type of water (based on water quality) required for various units in the
refinery
 Identification of problems in the water treatment system
Having gathered all the data with respect to water requirement of various units and the other water
requirement (service water drinking water, DM water, Fire water, cooling water, & condensate
recovery system) in the refinery first step towards a mass balance diagram for water is fulfilled.
The other information, which is required, is waste water generation and treatment facilities. The
data is collected with respect to effluent generated from different units in the refinery, effluent from
other sources, effluents quality and quantity, effluent collection, segregation and routing system. It
is important to find out the performance of existing effluent treatment facilities with respect to their
effectiveness & problem identification in existing ETPs to achieve Minimal national standards
(MINAS). After getting all information / data on the present practices and usages/ discharge of
water a block diagram is prepared wherein details of internal recycle/ treated water recycle/ reuse
of water get reflected.
Area of concern and action plan
The water balance thus prepared give a clear indication about the water guzzlers in a refinery. The
existing water management system is now critically reviewed with respect to effluent generation
from production plant operations, compliance requirement, environment related record /
documentation with respect to quality and quantity and collection philosophy of the effluents,
effluent treatment plant operation, etc. This is sometimes further substantiated by actual sampling
and analysis at inlet / outlet of all pollution control facilities to be identified after careful analysis of
water balance diagram. On the basis of all the collected data an assessment of the existing waste
water management system is to be carried out.
In order to overcome various problems like inadequate or less efficient treatment facilities or any
other reason like improper segregation and collection system, necessary corrective actions are
suggested in specific areas which may cover modifications / alterations / augmentation of existing
water and waste water treatment systems for the entire Refinery.
As a part of suggesting effective action plan effluent recycle and reuse alternatives, existing
refinery effluent disposal philosophy, optimization of raw water usage in the refinery, optimization
of existing refinery water treatment systems with respect to minimization of effluent and
maximization of effluent reuse, various alternatives for treating effluent from water treatment
facilities, optimization of collection, segregation routing and treatment of refinery effluents are
addressed. In such cases, the role of Zero Liquid discharge (ZLD) in meeting the legislative norms
and optimizing water requirement and formulating the required technologies to meet the same
become paramount.
Role of technology
In order to implement effective ZLD scheme, it is important to know how such plan can be
achieved, what are the latest technologies which can be implemented to achieve the effective
water optimization schemes.
Refineries have implemented technology such as low temperature multi effect distillation for sea
water desalination and mechanical vapour compression in such cases where sea water quality is
quite bad, Electro-dialysis to meet the required TDS at the outlet for discharge, ultra-filtration for
raw water treatment, ultra-filtration for DM water treatment for colloidal silica removal and as a pre
treatment step before reverse osmosis to control silt density index & reverse Osmosis for
desalination of waste water & sea water etc for the total water management. Similarly disk filtration
and micron filtration are some of the new filtration devices, which improve the efficiency of filtration
and utilizes fraction of water for cake cleaning. All the above technologies have been successful in
meeting the most stringent water standards & in reducing the overall water requirement of a
refinery.
Other technologies which are gaining momentum is the use of no chemical treatment in cooling
water circuit, as cooling water is the main water guzzler in a refinery. These technology not only
increases the cycle of concentration (COC) thereby improving the water quality. In above context,
magnetic treatment, pulse power technology and Galvanic cell technology are gaining importance
as they require less bleed off and less fresh water requirement.
Magnetic treatment: Allows minerals to precipitate out of solution in a less harmful fashion,
creating a softer, amorphous mud-like substance, known as aragonite, instead of calcite
Pulse power technology: It impart a broad spectrum of electro-magnetic fields into the flowing
cooling tower water and removes the static charge from naturally-occurring particles which results
in:
Scale Prevention: With the static charge removed the particles then become the preferred surface
for minerals to precipitate around rather than forming scale on the surfaces of the cooling tower
and HVAC equipment
Corrosion Control: The system operates in a saturated environment which is naturally noncorrosive. Additionally, there is no risk of chemical corrosion
Biological Control: Superior biological control is achieved through encapsulation and
electroporation. Dolphin treated water typically has even lower biological counts than the accepted
standards for drinking water
Galvanic Cell Technology: It exploits solubility characteristics of Calcium and Magnesium salts
in fluid with change in its pH value and locally increases pH value of water/fluid before it reaches
high temperature zone and then precipitates out hardness causing salts as water/fluid flows
through unit. It is an online pipe shaped mechanical device which houses very specially designed
core that is Sintered with number of electronegative elements in their order of increasing
electronegativity in the direction of water/fluid flow. As feed water/fluid passes through this unit, it
acts as an electrolyte while it comes in contact with the core placed inside it . Due to the Galvanic
action, the core gets negatively charged and attracts the lightest ions present in the fluid i.e
Hydrogen ion towards itself that decreases the pH of water and Ca and Mg salts gets precipitated
& does not affect heating surfaces. Consequently, once the precipitation of these hardness
causing salts has taken place, further precipitation due to temperature variation is not possible.
For oil removal the new technologies have targeted chemical bonding, reusable micron filtration,
and media adsorption and coalesce as techniques to arrive as a niche product, which can provide
answer to the present & future requirement of waste water treatment.
One such example is the use of Advanced Coalescer system which removes bulk oil and recovers
high purity skim oil and settles out large solids. The secondary regen system uses a media depth
bed to remove remaining bulk and dispersed oils and captures 98% of suspended solids >5
microns. Periodic regeneration of the media bed is performed to maintain high efficiency. The
Polisher system removes emulsified oils and fine solids to meet condensate water and cooling
water requirement < 1ppm of oil.
Water sustainability assessment
Major effluent streams generated in a refinery includes process effluent, spent caustic, DMP/CPU
regeneration wastes, and huge quantities of the cooling water blow downs. These effluent streams
are needed to be properly segregated, collected and then suitably treated in a effluent-cum-recycle
treatment plant to allow potential recycle of these streams within the complex. Some of the
effluents streams (e.g., spent caustic, DMP/CPU regeneration wastes, cooling water blow downs,
etc.) contain high levels of total dissolved solids (TDS) apart from other biodegradable & nonbiodegradable contaminants. While aiming at Zero Liquid Discharge, one should not forget the
Principle of Conservation of Mass, which makes disposal of these salts either in dissolved form or
in solid form extremely difficult or economically unviable especially for landlocked complexes. To
meet the objective of zero liquid discharge from the complex, a part of the water is recycled for
being utilized in the green belt around the complex, which requires control of TDS within
acceptable limits for green belt development. Whichever treatment option is employed, its
suitability in meeting the objective of zero liquid discharge is dependent primarily on the quantity &
quality of effluent generated and potential end users of treated water. In view of this, a careful
study of the overall water balance for the complex in view of the prevalent statutory environmental
norms, and zeroing down on the final water balance & methodology, which does not only meets
the ZLD objective but is also financially viable, is of uttermost importance.
The concept of pinch technology which has been successfully applied for heat exchanger
network has also been extended to water in a refinery. The application is analogous to the use of
Pinch Technology for minimizing energy demand .It has two approaches. Focus will be on the
interaction between waste water minimization and effluent treatment system design to minimize
cost of fresh water and treatment systems. One is waste water minimization and other is effluent
treatment system design. Various methods developed for minimization of fresh water and waste
water, using water pinch technology are in vogue in a modern refinery.
Abridged finding of some real cases of water assessment: In the context of Indian refinery
where various sources such as sea cooling water network, Raw water reservoir, Bearing cooling
water system, condensate reuse are the pertinent places which must be assessed carefully for
sustainability of water management in the refinery. Beside, refinery treated waste water provide of
the opportunity which should be utilized in order to bring down fresh water consumption in
processing units.
In some cases it was noticed that refineries where sea water cooling were practiced and there was
no filtration and sea water quality was bad in terms of oil & grease, BOD, COD & Suspended
solids, continuous bleeding was done to maintain temperature gradient across coolers & at times
back washing was performed to flush out settled silt at the coolers. The intentional water loss thus
noticed was to the tune of 25% of the makeup water resulting in energy and unsustainable
operation. The quickest solution to the problem was installation of on line filtration system.
Large open area in raw water reservoir sometimes provides more evaporation rate & loss of water.
This loss could go as high as 3%-4% which is significant. Beside heavy losses were reported due
to leakages of old piping network in the refinery which was to the tune of 10%.
In many cases bearing cooling water is not recovered & finds its way to ETP. Enormous saving is
achieved if a new network is created for the collection of water and it is returned back to the circuit.
In the context of Indian refineries the condensate recoveries are as high as 50%-60%, wherein in
any modern refinery condensate recoveries could go upto 80%. The recovered condensate can be
stored in a condensate recovery tank which can have multiple utilizations such as: cooling water
make up, treated filtered water and to steam plant.
Implementing a condensate recovery and re-use system will result in significant benefits to the
operations of the Utility unit which includes reduction of:
 Water charges due to reduced water demand
 Fuel costs as water will be reused in place of BFW
 Blowdown from steam generators and boilers
 Energy losses
 Chemical treatment costs
Usages of treated sewage for cooling water makeup water/ raw water quality is a better option
rather than using sea water desalination which is rather energy intensive, generates more effluents
and has high operating cost.
Conclusion
Refinery fresh water consumption today is guided by many factors that may relates to energy
and water efficiency, quality of water, internal recycle and reuse of condensate, efficiency of
wastewater treatment system, treated waste water quality to be effectively recycle and reused,
type of cooling water treatment system etc. For the sustenance of water management system
in a refinery, the role of water footprint can’t be ignored.
Condensate systems traditionally focused on the removal of oil from water prior to discharge
to drains and oily water sewers. Advances in technology, upgrades of process and plants and
stringent environmental regulations have resulted in systems where contamination by
hydrocarbon is an abnormal occurrence. The generation of clean condensate and its recovery
is one means to improve energy efficiency as well as benefit from cost savings due to the
reduced water consumptions via reuse, even if it is not recovered directly into a Boiler Feed
Water system.
Similarly, enhanced cycle of concentration using water integration and non chemical devices
or combination of both will result in saving lot of blowdown requirement. However it must be
kept in mind that increasing the COC will result in higher TDS of cooling water blowdown
which will consume high energy if cooling water is to be recovered through reverse osmosis
process. An optimal TDS balance is needed to be worked out for the entire water balance
circuit.
Avoid large open area for the storage of raw water, instead use deep tanks. Raw water treatment
plant if it is membrane based should be built up within the reservoir. Regular health checkup of
water line is must. Now a days there are better piping materials available, such as PE, GRP, and
CPVC etc which have been used in underground piping system, should be preferred which have
less problem of leakages.
Process & Reliability improvement in Reverse Osmosis (RO) system by
Innovative Chemical Dosing Modification
Mahendra Singh Bhadauria, Rajesh Nandanwar and Anand Pratap Raghav
(Bharat Oman Refineries Ltd, Bina)
Abstract:
The effectiveness and quality of Reverse Osmosis (RO) operations is a lot dependent on the
diligence of the right amount of chemicals dosed in the pretreatment system. A slight
overdosing or carryover of residual chemicals can hamper the RO operations severely.
Merely, robust pre-treatment systems like Ultra-filtration, Micron Filtration and Membrane
Bioreactors, are not adequate to mitigate the fouling issue in the RO system, simultaneously
chemical residuals needs to be appropriately guarded.
Membrane in itself is very complicated in terms of fouling mechanism. Effect and impact of
residual chemicals on membrane fouling cannot be accurately gauged. However, based on
good operating practices and regular watch on the operating parameters, chemical
consumption pattern, physical interaction between membrane foulants, water constituents,
and chemicals, it can be avoided further complicating an existing foulants problem.
Reverse Osmosis plant at BORL faced the issue of Antiscalant deactivation and chemicals
precipitation on the membranes and thus increase in pressure drop. This has resulting in
very high chemical cleaning frequency. Being a very serious issue in terms of sustainable
performance, detailed root cause analysis has carried out.
This has resulted not only in reducing the pressure drop issue but also helps in minimizing
the cartridge filter and chemicals consumption cost, thereby reducing the OPEX of RO Plant.
This paper provide details of methodology adopted in resolving the Antiscalant deactivation,
modification done in the system and remedial actions taken to improvise and sustain RO
plant efficiency and Reliability.
1. Introduction to Reverse Osmosis based De-Mineralization plant:
Reverse Osmosis based De-Mineralization (RODM) plant is provided to produce DM water at
the rate of 450m3/h on continuous basis to meet the demand of Refinery and Captive
Power Plant. DM water is used as boiler feed water make-up for the generation of steam
and as process water for dilution, reaction and washing.
RO system in the refinery is designed to treat treated effluent from ETP, treated water from
RWTP and waste water from cooling tower and boiler blow-down. These streams are being
used as feed in RO system for producing DM water.
2. Importance of Antiscalant in RO:
Polyamide membranes rejecting surface requires a better means of scale control. In the
absence of a softening process, there are several methods of inhibiting scale formation such
as pH control (suitable for calcium carbonate only), Antiscalant like sodium hexa meta
phosphate (SHMP) and a wide range of proprietary products. Today, the most widely used
scale inhibitors are based on an extensive range of organic compounds. These work by three
closely related mechanisms that interfere with one or more of the stages of crystal growth:
1. Threshold Effect — these inhibitors retard the precipitation of salts that have
exceeded their solubility products, e.g., phosphonate - based products.
2. Crystal Distortion Effect — These inhibitors distort normal crystal growth and
produce an irregular crystal structure with poor scale forming ability, e.g., polyacrylic
acid [CH2CHCOOH]n with molecular weights in the 1,500–2,500 range.
3. Dispersancy — Dispersants work by placing a surface charge on the crystal.
Phosphonate - based Antiscalants are excellent inhibitors for a wide range of scaling species.
I
e ra e syste s, they a t as super-threshold age ts. These produ ts have the a ility
to hold highly supersaturated solutions in a stable condition during the finite time it takes
the water to exit the membrane system.
3.0 Test & Analysis performed to identify foulants before discovering
Antiscalant Deactivation issue:
Following test & analysis being performed to identify membrane foulants
 Membrane Autopsy test carried out by Membrane manufacturer to identify the
nature of foulants deposited on the membrane surface.
 Membrane Autopsy test reported the presence of Aluminium on the membrane
surface and also the scaling due to silica.
 Silt Density Index (SDI) test carried on site and gum like precipitants were observed
on the filter paper.
 Deposit analysis test carried out to identify the chemicals residual material.
4.0 Symptoms of Foulants and their impacts:
Gum like precipitate presents on the membrane surface which was heavily fouling the lead
elements because of the chemical reaction of polymeric organic anti-scalants with
multivalent cations like Aluminium or cationic polymeric flocculants.
The source of silt or colloids in reverse osmosis feed waters is varied and often includes
bacteria, clay, colloidal silica and iron corrosion products. Pretreatment chemicals used in a
larifier su h as alu , ferri hloride, or atio i polyele trolyte s a also ause olloidal
fouling if not removed in the clarifier or through proper media filtration. In addition, cationic
polymers may co precipitate with negatively charged Antiscalant and foul the membrane.
5.0 Mitigation measures performed before discovering Antiscalant
Deactivation issue:
Following improvement measures taken to mitigate fouling issue in the RODM plant, before
discovery of Antiscalant Deactivation issue;
 At Raw Water Treatment Plant, Alum replaced with the advanced coagulant Poly
Aluminium Chloride (PAC) which minimizes the residual Aluminium carryover issue.
 Cationic polymeric flocuulants replaced with anionic polyelectrolyte in the pretreatment section of cooling tower and boiler blow down stream.
 Organic Antiscalant replaced with phosphonate based Antiscalant
 Pre-treatment facilities like Ultrafiltration and MBR checked for SDI performance.
Granular Activated Carbon filters were checked for COD removal performance and
found inline.
 Cartridge filter were tested with a range of better options available in the market.
 Overdosing and carryover of chemicals (like coagulants, polyelectrolytes, Dechlorinating agents, etc.) being controlled by performing the JAR test, before fixing
the dosage rates.
Slight plant performance improved like reduction in membrane cleaning frequency (from
2.5 days to 7 days), after implementing above improvement measures.
6.0 Methodology adopted in finding the Root Cause:
The best available technology for determining the fouling potential in reverse osmosis feed
water is the measurement of the Silt Density Index (SDI), sometimes referred to as the
Fouling Index (FI). This is an important measurement being carried out prior to designing a
RO pretreatment system and on a regular basis during RO operation (three times a day is a
recommended frequency for surface waters). We followed the same frequency and perform
SDI test in every shift at RO-I tank outlet, ETP outlet and ACF outlet.
We measured the SDI at all suspected locations to identify the root cause for high DP in RO
plant. Results of SDI are tabulated below table
Table No. 1: SDI Data
Location
MBR Outlet
SDI
(Design)
<3.0
ACF Outlet
RO- I Feed Tank
<3.0
<3.0
Cartridge
Outlet
Filter <3.0
SDI
Findings / Observations
< 3.5
SDI is high due to fibre damage in MBR,
rectified after repairing fibre
< 3.4
Slightly higher
< 3.3
Slightly on higher side but no major
impact
Not
Paper gets chocked due to chemical
Measurable precipitates.
First, symptoms of high DP problem being noticed when the SDI test even after many
repetitions will get failed due to clogging of filter papers. Clogging of 0.45 micron filter paper
and chemical precipitates indicated a clue that SDI might be increasing due to this and
creating a severe impact on Antiscalant efficacy.
Secondly, to freeze out the root cause and confirming the clue of chemical precipitation,
deposit sample was collected and sent to the Lab for detailed analysis. Foulants or deposits
can also be seen from the Fig. 1 below. Lab results also confirmed the presence of scalants
likes CaSO4, BaSO4, etc. and higher acid insoluble also reported more than 20%.
All chemical dosing points (HCL, Antiscalant and SMBS) were reviewed for finding suspected
locations for these chemical precipitation on cartridge filters. After reviewing this, some
shocking observations were noted as below

Both HCL & SMBS chemicals are reducing or degrading the Antiscalant efficacy and
becomes chemically inactive. HCL dosing point found very close to Antiscalant dosing
point (10 cm only)
Fig. 1 Cartridge Filter photograph with chemical precipitates
7.0 Remedial measures with innovative chemical dosing arrangement
After analysis of the lab results of deposit samples (from cartridge filters) and on-site results
of SDI test, following major innovations were discovered for improvements in the chemical
dosing points; which were rectified and elaborated as below

All the three chemical dosing points (HCL, Antiscalant and SMBS) are provided very
close to each other (with in 30 cm distance). Sufficient distance was not provided for
reaction time and solubility. Intermixing between chemicals was not considered in
the plant design.

Acid (HCL) dosing point, located very close to Antiscalant which tear down the
Antiscalant efficacy because of the intermixing with acids concentrated pockets.
Membranes including the cartridge filters become heavily fouled due to the
formation of these by-products or chemical precipitates. Chemical cleaning was not
effective even after thorough and repetitive cleaning with the specialty chemicals. 1st
stage membranes become heavier due to accumulation of scalants on the
membrane surface and clogging of pores.

Antiscalant dosing point provided at the upstream of Sodium Meta Bi Sulphite
(SMBS). SMBS dosing point is provided for removing residual chlorine or any other
oxidants from the RO feed water. Antiscalant chemical after reaction with residual
chlorine or any oxidants becomes deactivating. Slight variation in dosage rate was
actually hampering the RO plant operations. Antiscalant dosing point relocated after
cartridge filter at the downstream of the SMBS injection point to avoid a y eat
product mixing.

New Antiscalant dosing point has been provided at the downstream of cartridge
filters to prevent further deactivation of Antiscalant by iron.
7.1 Existing Chemical dosing Arrangement:
GAC Filter
HCL Antiscalant
SMBS
5 µ Cartridge
Filter Assembly
Fig-2 Existing chemical dosing arrangements
7.2 Chemistry Involved
•
HCL maintains RO feed pH <7.5, controls carbonate scaling
RO - I
•
SMBS Sodium Meta Bi-sulphite (De-chlorinating agent). Even very low levels of
chlorine or oxidants in the feed stream will result in irreparable oxidation damage of
the membrane.
7.3 Innovative Chemical Dosing Arrangement:
After performing various trial runs with this successful and innovative design, it was decided
to relocate HCL and Antiscalant chemical dosing points based on modified chemical dosing
arrangement as per below scheme
HCL
GAC Filter s
SMBS
Antiscalant
RO - I
5 µ Cartridge
Filter Assembly
Fig. 3 Innovative Chemical Dosing Arrangement
7.4 Major benefits after innovative modification:
•
Chemical cleaning frequency reduced from 7 days to 76 days
•
Cartridge Filter replacement frequency improved from 7 days to 45 days
•
Higher membrane flux improves the recovery of RO Plant which was earlier
deteriorating because of frequent chemical cleanings and clogging of membrane
pores.
•
Chemical consumption of Antiscalant, SMBS, HCL, Cleaning chemicals reduced.
•
Cartridge Filter consumption reduced can also be seen from the table below
Table No. 2: Comparison of RO Plant performance after innovative
modification
Parameters
UOM
Chemical Cleaning
days
Frequency
Cartridge
Filter
days
Replacement
Antiscalant
Kg / MT of
DM Water
Design
Before
After
Improvements
in %
90
7
76
985
60
7
45
543
9.10
6.52
5.2
20
Hydrochloric Acid Kg / MT of
(HCL)
DM Water
Non-Oxidizing
Kg / MT of
Biocides
DM Water
Kg / MT of
High pH Cleaner
DM Water
Kg / MT of
Low pH Cleaner
DM Water
Overall Recovery of
%
RODM Plant
Will vary as
per feed pH
50.9
33
35
1.0
0.65
0.1
85
2.6
0.60
0.4
33
5.1
1.13
0.5
56
66
59
63.6
7.8
8.0 Conclusion:
A proper diagnosis with innovative approach helped us in discovering and mitigating the
high ΔP issue i RO pla t due to A tis ala t dea tivatio . I
ediate re edial easures
taken and chemical dosing point location modified as per in-house innovative chemical
dosing arrangements.
Barely, robust pre-treatment systems like Ultra-filtration, Micron Filtration and Membrane
Bioreactors, are not adequate to mitigate the fouling issue in the RO system, simultaneously
chemical residuals needs to be appropriately guarded. The complexities of membrane
fouling and the potential adverse effects of chemicals residuals cannot be ignored.
However, with a practical knowledge of these chemicals and their physical interactions
between water constituents, foulants, and chemical additives, it can be avoided further
complicating an existing foulants problem.
Author: Raju Chopra, Maninder Jit Singh, Sagar Shukla, Sachin Panwar
Haldor Topsoe India
Title: “Energy efficiency improvement through innovation”
Introduction
India is the 3rd highest energy consumer and its average energy consumption growth is 5.3% which is
projected to grow further. India’s energy needs are expanding rapidly. As India being highly dependent
upon crude export, it becomes more important that we become more and more energy efficient. As
there is growing emission concerns worldwide, it becomes extremely important to focus on the energy
efficiency improvement in the process technologies in refinery and petrochemical industries. Topsoe
has been continuously endeavoring in technology and catalyst development area in order to offer most
energy efficient solutions to their customers. Our strong R&D capabilities, in-depth knowledge about
reaction kinetics, sound engineering experience and extensive portfolio allow us to offer the most
energy efficient solution to our customers. Due to innovative approach towards energy efficiency
improvement, Topsoe is in an unique position to offer various technology modules to refinery and
petrochemical industry. This paper covers the various techniques and technologies offered by Topsoe
to improve the refinery overall energy efficiency mainly in hydroprocessing, hydrogen generation and
sulfur management areas.
Hydroprocessing
Use of latest generation catalyst system
One of the key element in hydroprocessing units is high performance new generation catalysts which
allow refiners to operate unit at much moderate conditions w.r.t. temperature and pressure saving
substantial amount of energy.
Topsoe’s CoMo BRIM® catalyst series allow refiners to even produce the ultra-low sulfur diesel at low
reactor pressure as 25-30 bar g while optimizing the hydrogen consumption. There are many refiners
worldwide who face hydrogen shortage during the refinery operations and wish to produce ultra-low
sulfur diesel. Topsoe allow such refiners to utilize the CoMo BRIM® catalyst series which follow direct
desulphurization route to meet the diesel sulfur while minimizing the hydrogen consumption. High
activity of catalyst can be utilized in many ways i.e. lowering the unit pressure, running the reactor
charge heater at lower temperature, lowering the recycle gas rate which improve the energy efficiency
of the unit.
In this competitive world, every refiner is willing to install a new unit or revamp existing unit with low
capex and opex. Catalyst selection is a very important step during unit design. Units where objective is
primarily desulphurization, Topsoe CoMo BRIM® (TK-568, TK-570, TK-578) catalyst series can be
utilized which can operate relatively much low reactor pressure starting at 25-30 bar g. Similarly, for
units targeting large improvement in density, cetane improvement; Topsoe NiMo HyBRIMTM (TK-565,
TK-569, TK-609, TK-611) catalyst series can be utilized which can operate relatively at much low
reactor pressure starting at 45-50 bar g. Low pressure unit operation with high performance catalyst
reduce the unit specific energy consumption to a large extent allowing refiners to improve their gross
margin.
Similarly, Topsoe 2nd generation NiMo HyBRIMTM catalyst series and red/blue series of hydrocracking
catalysts also allow refiners to run the VGO HDT, Mild hydrocracker and Hydrocracker units at
relatively lower reactor pressure and temperature resulting into substantial reduction of capex and
opex.
Graded bed solution
In the trickle bed reactor, right type of grading to be ensured to avoid the pressure drop issue. Two
distinct mechanisms affect the way in which the void space is reduced or plugged. 1) Deposits on the
catalyst bed of solid contaminants carried over with the feed, 2) Deposits on the catalyst from reaction
products originating from components in the feed. The premature pressure drop across the reactor
bed, leads to less energy efficient operation where recycle gas compressor has to compress more and
as a result of high pressure drop, more power is needed by feed pump and make-up hydrogen
compressor. Topsoe offers special grading TK-10, TK-15 and TK-26 Top trap which contains much
larger void volumes to mitigate pressure drop issues and ensure energy efficient operation throughout
the guaranteed cycle length.
Process design optimization
Hot vs Cold separator layout
During the unit configuration, Hot and Cold separator layouts are evaluated to choose the best in terms
of energy efficiency.
Advantage of hot separator lay-out
It requires lower investment cost which can be lower by ~40 US$/bpsd at pure make-up gas. Hot
separator layout requires:
-
Fewer heat exchangers
-
Smaller reactor effluent air cooler
-
Lower head compressor
-
Smaller wash water flow
This layout requires low energy consumption up to ~9500 MJ/bpsd/year which results into lower
heating requirement.
Figure 1, Hot separator layout
Disadvantage of hot separator lay-out
-
In this layout (refer Figure-1), there is increase in impurity accumulation which reduces the
hydrogen partial pressure. It required more catalyst volume and/or higher unit pressure to meet
the processing objectives. This layout becomes uneconomical when total influx of methane
becomes higher than 3-5 Nm3/m3 (typical), unless very large recycle gas purge (refer Figure-2).
-
This layout leads to higher H2S accumulation, for example for a feed with 1.4% sulfur,
@ 75 bar g, Cold separator gas contains 3 vol% of H2S treat gas whereas Hot separator gas
contains 10 vol % H2S treat gas which require amine scrubbing to remove the H2S from gas.
-
In this layout, +2-3 Nm3/m3 hydrogen is dissolved in the hot separator liquid thus leading to
higher overall hydrogen consumption.
Figure 2
Addition of LP flash drum
In medium to high pressure hydrotreaters, LP pre-flash drum can be installed to recover the dissolved
hydrogen by flashing of high pressure hydrocarbon liquid and hydrogen rich gas at ~20-24 bar g can
be routed for purification in PSA or re-use in other hydrotreaters within the refinery complex. A
hydrogen saving potential of ~3 to 12 Nm3/m3 is expected in hot separator configurations depending
upon pressure level.
Heat integration in hydrocracker unit
The hydrocracker poses an interesting design challenge for safe and effective heat integration
between the high heat release reaction section and the high energy demand fractionation section (refer
Figure-3). Topsoe recommends adding a reactor effluent/fractionator feed heat exchanger in
hydrocracker unit design. Our studies have shown that by adding this heat exchanger can reduce the
fractionation feed heater duty by 1/3 to as much as 2/3, depending on the battery-limit raw feed
temperature.
Figure 3
Utilization of coker off-gas to produce hydrogen
Utilizing the Topsoe Fuel Gas Hydrotreating technology (refer Figure-4), coker off-gas can be treated
to remove the sulfur, COS, CS2 and saturate the di-olefin and olefins. Treated gas can be used as
feed to hydrogen generation unit replacing the more expensive feed like naphtha and natural gas.
Figure 4
Optimization of recycle gas to feed ratio
For a given feed rate, the cost of a hydrocracker is strongly influenced by the reactor pressure and the
hydrogen circulation rate. The unit pressure is typically set to obtain the hydrogen partial pressure
required to meet the product quality and to minimize catalyst deactivation rate. Topsoe has the studied
the effects of gas to oil ratio on hydrogen partial pressure, cracking reactor temperature rise, catalyst
activity, and stability and concluded that we can reduce the gas to oil ratio of a new hydrocracker
without sacrificing performance and safety. Optimum recycle gas to feed ratios can save substantial
amount of energy due to reduction in HP loop pressure drop and reactor pressure.
Various other ways
There are various other ways to improve the energy efficiency:
-
Using stepless control for reciprocating compressor
Using the power recover turbines for HP liquids (HP liquid from HPHS, HP amine absorber)
Use of variable frequency drives for feed pumps, HP amine pumps, RGC
Using special type of equipment i.e. Helical baffles, twisted tubes, Packinox exchanger
(naphtha, kerosene and light diesel services)
Optimal split of air cooler and water cooler duty and usage of variable air cooler fan blade pitch
Common overhead system for stripper/stabilizer section
Recovery of hydrogen from purge gas and recycle recovered hydrogen to make-up hydrogen
compressor
Hydrogen generation
Steam reforming is the dominant and reliable technology for producing hydrogen or hydrogen-rich
synthesis gas. Feedstock for producing hydrogen or hydrogen-rich synthesis gas may range from
natural gas to fuel oil/ Naphtha. Over the past 50 years, Topsoe has continuously focused on the
development of new steam reforming technologies resulting in the design of more than 250 steam
reforming units.
In the pursuit of improved energy efficiency and more flexible solutions for hydrogen production
Topsoe has continued to develop and improve technologies for heat exchange reforming since 1985,
when the first Heat Exchange Reformer was released. Today, more than 30 years of development and
operating feedback, Topsoe has generated a unique knowledge and experience in the field of heat
exchange reforming. Haldor Topsoe Exchange Reformer (HTER) was developed in the late 1990’s and
demonstrated in full-scale from the beginning of 2003.
In the traditional synthesis gas plants, the sensible heat in the synthesis gas has been used for steam
generation. The sensible heat in the gas from the reformer can be utilised better, as the temperature
difference in the hot end typically is in the range from 650-750°C. However, due to the corrosion
phenomenon known as metal dusting, it is a challenge to design the heat exchange apparatus in such
a way.
Being an endothermic reaction steam reforming of hydrocarbons require a significant heat input to
obtain the desired conversion into hydrogen. In the conventional steam reformer heat transfer takes
place by radiation which leads to a limited thermal efficiency as evidenced by a high flue gas
temperature. The thermal efficiency of a conventional steam reformer is around 50% and the surplus
heat is used for steam production. Many refiners have little or no use of the steam export generated in
a hydrogen plant which is therefore considered of low value.
Heat exchange reformers are characterised by being very compact and by a high thermal efficiency. In
a heat exchange reformer, the majority of the heat transfer takes place by convection with hot flue gas
or hot process gas whereby the thermal efficiency can be increased by as much as 60-80% as
compared to the radiant solution. Conceptually, a heat exchange reformer uses the waste heat energy
(carried by process or flue gases) to produce extra hydrogen instead of surplus steam.
Figure 5: HTER in H2 Plant
Haldor Topsoe Exchange Reformer (HTER)
The Haldor Topsoe Exchange Reformer (HTER) is a process gas heated reforming concept in which
the hot reformed gas, at high pressure and temperature, is used as heating medium. The high
pressure enables a more effective convective heat transfer. The HTER is well suited for capacity
revamps and new units in which factors such as efficiency, compactness, load following capacity etc.
are important.
The HTER can be used as a parallel concept as seen in Figure 7.
In the parallel HTER concept the feed gas is split in two streams. One stream is sent to the main
reformer where it is heated and converted. The other stream is sent to the gas heated reformer. The
main part of the heat for the gas heated reformer is provided by cooling of the effluent from the main
reformer. A part of the reaction heat will, as for the series HTER, be provided by heat exchange of the
Feed Gas
Product
gas
Reformed gas
Figure 6: HTER concept
Figure 7: parallel HTER concept
reformed gas with the catalyst bed.
HTER Benefits
HTER is a compact and efficient technology which provides unique revamp opportunities. Benefits
include








Minimum modification of existing plant
Easy operation
High efficiency
+10-30% reforming capacity
Low investment
Very small plot area
3-6 weeks unit down time for revamp
Lower CO2 emission
Case stories

Figure 8: Essar Oil HTER
Grass root project : ESSAR, India
In collaboration with Essar Oil, Topsoe evaluated different options and it was found that Haldor
Topsoe Exchange Reformer was the most feasible solution which utilizes the sensible heat in the
reformed gas for additional reforming instead of steam production. ESSAR oil was a Grassroot
hydrogen project and it was commissioned in 2012 with Double tube HTER-p. The benefits of the
implementation of the HTER in Essar’s plant is seen and compared to a similar plant without HTER
in Table 1.
Table 1: ESSAR operating data

H2 plant w/o
HTER
H2 plant with
HTER
H2 production
Nm3/h
130,000
130,000
Feed + Fuel
consumption
Nm3/h
55,366
51,485
Steam export
t/h
103
40
Capacity enhancement project: NRL, India
NRL was a hydrogen revamp project and it was commissioned in
2010 with Bayonet tube HTER.
The timeline for revamp was short restricting the downtime of
refinery to just six weeks. Below is a diagram shows the timeline
briefly
Figure 9: NRL HTER
Shut down
HTER installed
4 weeks
Start-up
2 weeks
Figure 10: HTER revamp timeline
The benefits of the implementation of the HTER in NRL’s plant is seen and compared in Table 2.
Table 2: NRL operating data
Before
revamp
After
revamp
H2 production
Nm3/h(Δ%)
Feed + Fuel consumption
Gcal/1,000 Nm3 H2
3.63
3.39 (-6.5%)
Steam export
t/h
34.1
16.3
52,816
67,419 (+28%)
.
The innovative heat exchange reforming solutions developed and industrially demonstrated by Topsoe
are the optimal solutions for boosting the capacity in existing hydrogen units. The flexibility,
performance, cost, implementation time, etc. are unmatched by other revamp options. The diversity of
the Topsoe exchange reforming technologies allows for optimal utilisation of existing plant capacity in
case of revamps and optimized production cost for new plants.
For new plants heat exchange reforming offers a competitive option for minimising the production cost
of hydrogen and is an environmental attractive solution due to reduced emissions.
The options offered by heat exchange reforming for optimising the layout of new plants and utilising the
full revamping potential of existing units is to be considered when hydrogen capacity increase is on the
agenda.
Sulfur management using Wet gas sulfuric acid (WSA) technology
Oil refineries today are facing increasing challenges in managing the sulfur generated when
processing crude into marketable products. There is an increasing number of refiners that are now
processing cheap sour crude with high sulfur content. This combined with the stricter sulfur limits in the
refined products gives a significant amount of sulfur residuals which needs to be handled one way or
the other. At the same time, gaseous emissions from the refineries are constantly under pressure as
the legislation gets increasingly tightened. All these factors make the choice of method for managing
the sulfur very important and there can be significant potential savings by making the right choice.
Conventionally in refineries and petrochemical industries, steam is generated by burning the
hydrocarbon (fuel oil, naphtha, fuel gas, NG). Topsoe offers a unique Wet Gas Sulfuric Acid (WSA)
technology solution which has large applications within oil refining and petrochemical industries is one
such example which convert waste streams like acid gas, sws gases, flues gases etc. into commercial
grade sulfuric acid with potential sulfur recovery of >99.9% with high energy efficiency resulting in large
export of superheated high pressure steam.
The WSA process is sensitive neither to hydrocarbons nor to nitrogen containing compounds such as
NH3 in the feed gas, and it is also able to accommodate significant changes in feed gas flow and
composition. This is mainly due to the fact that in the WSA process there is complete combustion as
initial step whereas the Claus process has a somewhat more complex combustion.
A WSA plant will typically have lower investment cost when compared with a Claus plant equipped with
downstream tail gas treatment unit (TGTU). This is mainly due to the simple WSA design consisting of
relatively few pieces of equipment. When considering operating costs, the WSA technology also has
some important benefits. The large amount of high pressure steam produced in the WSA process has
a positive impact on the operating income. Finally, the relatively simple design and operation of a WSA
plant results in lower operating and maintenance cost compared to the more complex Claus plant.
Higher energy efficiency - Going from H2S to H2SO4 instead of stopping at elemental S is much more
favourable because of thermodynamics. This means that due to the number of exothermic reactions
occurring in the WSA process, about four times the amount of steam, of higher quality, is produced
when compared to the Claus process. The difference in reaction enthalpies between claus and WSA,
illustrated in Figure 11, means that WSA plant exports much more energy as compared to Claus plant.
Figure-11
Figure-12: Typical WSA layout
Most of the heat released is converted to superheated high pressure steam at the cost of treating acid
gases.
The WSA consists of three steps. In the first step the gases are combusted to produce SO2. The flue
gas from the combustion is cooled to approx. 400°C in a waste heat boiler, generating high-pressure
steam at around 60 bar g. The second step is the exothermal conversion of SO2 to SO3. This takes
place in a three-bed catalytic SO2 converter loaded with Topsoe VK-W series catalyst. Cooling
between the beds takes place by superheating of the steam from the waste heat boiler to typically
400°C, 45 bar g. The cooling after the last bed to approx. 290°C takes place by additional steam
generation in a boiler connected to the same steam drum as the waste heat boiler. During this cooling
about half of the SO3 generated in the SO2 converter will react with water vapor in the gas and form
sulfuric acid vapor. This is like the SO2 conversion an exothermal process, and the reaction enthalpy is
recovered in the form of steam. The third step is the condensation of the sulfuric acid vapor. This takes
place in a heat exchanger with vertical glass tubes cooled on the outside by atmospheric air. This
“WSA condenser” is the heart of the process. Condensed sulfuric acid leaves the condenser and can
be cooled and pumped to storage. The gas that is now almost free of sulfur compounds can be
directed to the stack. The air that has been used to cool the condenser is available at around 240°C as
combustion air in the combustor. In this way all the heat of the gas down to 100°C is utilized for steam
production, refer Figure-12 for typical process scheme.
Moreover, refiners using petroleum coke and/or other heavy residues for production of steam and
power will enjoy benefits from having a rather cheap fuel source for the steam/power generation,
offering an attractive alternative use of the bottom of the barrel. However, one drawback of burning
these heavy residues in a utility boiler is that flue gases generated by combustion contain high levels of
SO2 and NOx as a result of the high sulfur and nitrogen content in the fuel. These pollutants need to be
removed from the flue gas before being released to the atmosphere. The main challenge with flue gas
desulfurisation (FGD) conventional methods such as wet-FGD using limestone is that it gets
increasingly expensive and troublesome to operate as fuel sulfur content increases, i.e. higher flue gas
SO2 emission. For the above situation, the unique Topsoe SNOXTM technology is a superb option to
avoid excessive operating costs for the flue gas cleaning. The SNOX TM process efficiently removes
SOx, NOx and particulates from the flue gas without generation of waste and without consumption of
water, adsorbents and chemicals, except for a small consumption of ammonia used for NOx reduction
which can be replaced with sour water stripper gas if available. The product of the process is
commercial-grade sulfuric acid and energy produced from reactions in the SNOX TM plant is recycled
back to the boiler in the form of hot air, increasing overall boiler efficiency. In the below case study
(refer Table 3), utility consumption, total utility consumption cost and net production cost comparison
has been made between WSA and Claus + TGTU process, where WSA is producing commercial
grade sulphuric acid and Claus + TGTU process is producing elemental sulphur. It is quite apparent
from this comparison that total production value in case of WSA is much higher and total consumption
cost is much lower as compared to the Claus + TGTU which makes WSA technology much attractive
amongst other similar technologies.
Table 3
Basis
Feed Stream
Content of H2S in ARU
Total Sulfur Amount
Sulfur Recovery
Operating time
Plant life
Production
Figures
Sulfuric Acid (as 100%)
Sulfur (as 100%)
High Pressure Steam
Low Pressure Steam
Total production Values
NM3/hr
vol %
MTPD
%
hr/y
years
860
95
28
99.5
8600
20
Unit
Unit
price
INR
MT
MT
MT
MT
lacs INR
WSA
Claus + TGTU
Prod/Cons lacs INR Prod/Cons lacs INR
per hour
per year per hour
per year
3500
3.56
1071
10000
1.16
997
964
9.50
787
825
2.70
191
1858
1188
Consumption
Figures
Unit
Unit
price
INR
WSA
Claus + TGTU
Prod/Cons
per hour
Fuel Gas
Oxygen
High Presssure Steam
Boiler Feed Water
Cooling Water
Fresh Water
Electrical Power
Waste Water Treatment
Total Consumption cost
Total production Values
Net Production Income
MT
MT
MT
MT
MT
MT
kwh
MT
lacs INR
lacs INR
lacs INR
25600
2820
964
343
0.95
1.8
6.35
7.2
lacs INR Prod/Cons lacs INR
per year per hour
per year
0.23
506
0.56
136
1.33
110
11.20
330
3.70
109
30.80
3
1.00
0.15
205
112
400
218
1.00
1
445
1080
1858
1188
1413
108
Some of the other inherent advantages with the WSA technology are
-
high conversion of sulfur without the use of an expensive Tail Gas Treatment Unit (TGTU)
-
high flexibility towards fluctuations and components in the feed, easily processing nitrogen
containing compounds and hydrocarbons (incl. BTX)
-
simple and flexible operation with high turndown ratio
-
simple process scheme with a simple layout consisting of few pieces of equipment (i.e. low foot
print)
-
No water consumption and no solid waste generation
low pressure operation of WSA, allow Sour water stripper and Amine regeneration units to run
at lower pressure saving substantial amount of energy
WSA technology allows customers to meet the tougher environmental norms in most efficient manner
while improving their gross margin. Today, about 38 WSA plants are contracted in refineries worldwide
and in total more than 130 WSA plants are under construction or in operation.
Topsoe Furnace Manager (TFM)
Fired heaters in refinery processes pose numerous challenges. Among them are fire, process tube
failures, poor burner operation, mechanical equipment failures and process upsets. Historically, these
challenges required plant personnel to directly interface with the firebox to evaluate equipment and
process status. A critical activity, this interaction involves inherent risk. While specialized personal
protective equipment and hand-held data collectors have been developed, risks remain. Continuous
firebox monitoring technology relieves the dependence upon human firebox interaction. The Topsoe
Furnace Manager system is a strong, engineered, turnkey, administrative safeguard in operation
continuously monitoring the furnace firebox. It is in operation providing direct feedback with firebox
images, and providing interpretive data during all phases of plant operation – normal and abnormal –
throughout startup, shutdown, malfunction, maintenance and normal operations. Topsoe Furnace
Manager bridges the inevitable gap between engineering safeguards and administrative safeguards
utilizing state-of-the art technology with common, sensory feedback provided directly to the trained
operating staff continuously for performance & efficiency optimization. Topsoe Furnace Manager also
provides direct feedback to remote resources (i.e., experts) available to assist and interpret unusual
situations encountered throughout all phases of the furnace lifecycle.
Figure-13
Topsoe Furnace Manager module (refer Figure-13) can be applied to refinery and petrochemical
furnaces. The various advantages of installing a Topsoe Furnace Manager in furnace like Primary
reformer can be seen below:

Optimization of reformer firing increasing feed utilization by increasing temperature of the
reformed gas outlet the reformer.
The best performance of the reformer is obtained when it is operated at maximum/design outlet
temperature. Such optimization exercise is time consuming and requires quit a number of manual
TMT measurements. The Topsoe Furnace manager will read the TMT online making possible to
display the TMT temperatures in DCS system for optimization. Optimization helps in minimizing
methane content, increase of steam production, better performance of PSA and H2 production
enhancement. Thus for same production requirement feed and fuel can be optimized for energy
efficiency.

High Temperature protection of tubes
Topsoe’s experience is that premature tube ruptures in most cases are caused by overheating of
the reformer tubes during start-up and shut down. This can happen without direct warning to the
panel operators. Installing a Topsoe Furnace manager system will warn the panel operators if the
reformer tubes are heated beyond design. A reformer tube rupture will normally result in a 5 to 10
days unplanned stop. As a rule of thumb is that the impact on the revenue for a refinery stopping a
100.000 Nm3/hr hydrogen plant will be some 1 million US$/ day.

Increased safety in relation to burner ignition during start-up and shut down.
Ignition of burners always contains a risk. Today operators are relying on manual verification of
ignition. The camera system will, provided the right procedures are in place, increase safety as the
system automatically will generate a burner ignited feed back to the panel operator.

Remote access to reformer performance data.
TFM collects millions of data points annually, with deviation alarms stored in an easy to use
historian, and is remotely available to the organization outside of the plant. Experts can access
furnace information remotely, at any time, and provide timely feedback to frontline operations and
maintenance personnel. The remote viewing capability of TFM, supported by the capture of images
and data stored in the historian, enable communications, planning and organizational coordination
during periods of furnace problems and allow operator to optimize the overall performance of
furnace at various plant loads.

Enhanced Green foot-print
TFM helps in reduction of fuel as it helps in fast desired equi-distribution of heat and optimization in
furnace for respective operation load. This results in less generation of air pollutants like SOx, NOx
and CO2.
Economics benefits
There are several benefits that accrue from the installation of the Topsoe Furnace Manager:
Burner balancing and heat distribution
The system provides information to guide operations through the burner balancing process by
providing images and data before, during and after adjustments. With better heat distribution, fuel gas
consumption can improve 1% to 2% at the same production rate with savings up to $100,000 per year.
Burner balancing also promotes longer tube life by distributing heat flux more evenly. With ongoing fuel
savings and longer tube life, continuous fire box monitoring pays for itself between outages.
Sample payback calculation of an unplanned furnace incident resulting in a multiday outage:
Furnace processes
: 5,000 bbl/d
Refining margin
: $5/bbl
Labor and materials to repair furnace
: $50,000
Continuous monitoring system cost
: $400,000
Simple payback
: ($400,000-$50,000) / (5,000 bbl/d x $5/bbl) =14 days
Sample payback calculation of an unplanned furnace incident resulting in an outage of several
days:
Continuous monitoring system cost
: $400,000
Startup/shutdown costs
: $100,000
Repair costs
: $100,000
Per-day margin loss during operation
: $25,000
Payback
: $400,000 - $200,000 / ($25,000) = 8 days
Incident avoidance economics
Continuous firebox monitoring pays for itself in a few days by helping to avoid margin loss associated
with an unplanned outage, by avoiding repair costs and startup/shutdown costs associated with the
outage.
Topsoe Furnace Manager advantages can be concluded in ripple effect diagram in Figure-14.
Figure-14, Topsoe Furnace Manager Ripple effect diagram
Topsoe Support
The pre-engineered, standardized design of the Topsoe Furnace Manager system ensures that
adequate lifecycle support will be available 100% of the time. Topsoe is known for deep understanding
of processes, technology, catalysts, and ongoing support. The Topsoe Furnace Manager is supported
by Haldor Topsoe, with periodic support visits, troubleshooting, calibration, training, and consultation
readily available as part of a service agreement. With a deep understanding of chemical processes
and technology, and a very significant global network of expertise, Haldor Topsoe is the global leader
in integration of process support. Topsoe Furnace Manager is supported through this network, and
provides continuity as personnel transition through operating facilities. Haldor Topsoe working with
Topsoe Furnace Manager Technology can help bridge the gap when knowledge transfer is required
during periods when newer personnel become involved with furnace operation and maintenance. In
addition to the benefits of safety, reliability, collaboration and operability, Topsoe Furnace Manager is a
strong operational excellence platform to facilitate communications and knowledge sharing among the
entire organization. It is also important to note that the Topsoe Furnace Manager will enable operation
to easily balance to the firing evenly throughout the furnace resulting in reduced fuel consumption
which will result in significant OPEX saving and reduced in CO2 emissions during plant operation.
Topsoe has been offering various innovative ways to improve the energy efficiency of their offered
solutions which is apparent from the examples given in this paper.
Indian Oil Corporation Limited
Paradip Refinery
Technical Services
Maximization of BS-VI MS at Paradip Refinery
Geethashree, AM (PS), Rahul Srivastava, STSM & Amal K. Roy, DGM (TS)
1.0 Introduction
Across the world over the last many decades, atmospheric air quality has been adversely
impacted by emission from automobile tailpipe exhaust, industrial smoke stacks, thermal
power plants, construction dust & debris and the other by-products of a crowded and
modernized urban existence. Simultaneously the rising incidence of a range of health effects
has been recorded and there is compelling evidence of a causative link from the former to
the latter, some very direct, some somewhat direct and some in an associated sense along
with other factors.
It is true that deterioration in ambient air quality is not the sole source of stress on the lives
and health of our citizens. Nor is vehicular tailpipe emissions the only source of air borne
pollutants. However, vehicular emissions are indeed a large contributor to air borne
pollution which is the primary reason for mandating stringent fuel and emission standards
leading to shifting of fuel (gasoline and diesel) specifications from BS-IV (50 ppm sulphur) to
BS-VI (10 ppm Sulphur) standard.
2.0 Brief on Paradip Refinery
Paradip Refinery has a crude processing capacity of 15 MMTPA. The refinery initially
configured as a full refinery (refinery integrated with petrochemical complex) was reconfigured as fuels only refinery in Phase-I and subsequent integration of the refinery with
petrochemical complex (PP and PX-PTA) in Phase-II in order to reduce Project cost. This
deferment of Petrochemical complex had huge impact which led to change in mode of
operation of CCR from Aromatics mode to MS mode (due to deferment of PX-PTA) and FCC
INDMAX operation from Petro mode to MS mode (due to deferment of PP). This change in
refinery configuration had huge impact on the LPG and MS production figures. A
comparison of the MS, Naphtha and LPG yields are indicated below.
Attribute
Full Refinery
Fuel Refinery
Fuel Refinery + PP
kTPA
Crude
15000
15000
15000
Naphtha Sale
Gasoline (MS)
LPG
498
1353
784
87
4080
420
215
2076
896
This mega refinery was successfully commissioned progressively from Apr’1 to Nov’1 . The
refinery configuration consists of secondary processing units like DHDT and VGO-HDT for
production of BS-IV compliant diesel and FCC INDMAX, CCR and Alkylation units associated
with other tertiary units for production of BS-IV compliant gasoline. The deferred Poly-
Propylene plant is presently under construction with expected commissioning by 2018 and
hence FCC will be operated in Petro mode thereafter. However, due to deferment of PXPTA, CCR will continue to operate in MS mode.
Comparison of product yields: Design vs Actual
4080
kTPA
4500
4000
3500
3000
2500
2000
1500
1000
500
0
2076
1353
498
784
87
215
NAPHTHA SALE
GASOLINE (MS)
Full Refinery
896
420
LPG
Fuel Refinery
Fuel Refinery +PP
The MS blend for the refinery consists of Reformate from CCR, treated Light Naphtha from
FCC, Alkylate from Alkylation unit and Light Hydrotreated Naphtha ex NHT-NFU. As shown in
the figure above, the actual MS production in the operating scenario has decreased from
4080 TMTPA to 2076 TMTPA. This reduction in MS production is attributed to operation of
FCC in Petro mode where MS generation is lower by design, lower reformate absorption
into MS pool (due to higher aromatics content) and lower availability of alkylate (zero
aromatics and benzene content) from Alkylation unit.
Alkylate
Reformate
LT HDT Naphtha
HY HDT Naphtha
FCC LCN Treater
FCC MCN
BLEND
MS Euro-IV
Manufacture SPEC
Mass,
TMTPA
RON
% OLF vol
% AROM
vol
% Benzene
vol
Sulfur,
ppmw
377
890
313
20
341
135
2076
96
104
59
54.5
92
96
91.5
0.3
0.85
0
0
54.2
20.8
10.7
1
82
2
18.8
2.3
34.8
34.9
0.2
1
1.75
0
1.2
0.6
0.97
8
0.5
0.5
0.5
44
353
31.9
91.6 /
91.2
21%max
35% max
0.9% max
50
The original configuration envisaged CCR operation to create feedstock for PX Unit as well as
for production of Hydrogen to meet the refinery demand. However, change in operating
mode of CCR from Aromatics to MS has also resulted in operation at reduced severity thus
impacting Hydrogen production. Hence, in the present scenario, to meet the Hydrogen
demand of the refinery, CCR Unit is forced to operate at high capacity leading to generation
of excess Reformate, which is not getting absorbed in MS pool but is being sold as
Reformate.
Alkylation unit requires Butene (Olefin) and isobutane (Saturates) to produce high octane
alkylate. LPG ex PRU is the only olefinic rich stream with olefin content of ~60 – 70 wt%
along with saturates of 30 – 40 wt%. The design saturated streams are from AVU & CCR.
Alkylation unit is integrated with Butamer unit for isomerization of n-butane in the
saturated streams to iso-butane required for Alkylate production. CCR LPG being the major
contributor of the saturated LPG for Alkylation unit has resulted in shortage of saturated
LPG due to change in mode of operation of CCR. Hence, even after complete utilization of
the saturated LPG streams of AVU & CCR, Alkylation unit is under-utilized due to the
shortage of isobutane. Further, the utilization of saturates present in the PRU LPG is limited
by the olefin content of this stream.
3.0 Refinery Shifting from BS-IV to BS-VI quality fuels
With the advent of Auto fuel policy 2030 with requirement of upgradation of fuel quality
with respect to Sulphur, need arises for inclusion of additional facilities for lowering the
sulphur content of the fuels to less than 10 ppm, as the only specification that is changing in
BS-IV and BS-VI is the Sulphur content of the fuel.
In order to gear up refinery for the fuel quality upgradation, different scenarios have been
studied using LP model and paper blends. The cases studied for the purpose are 




Base Case: Existing refinery configuration (Fuel Refinery post PP with BS-IV MS)
Case-I: Existing configuration with maximum generation of BS-VI MS
Case-II: Installation of New Naphtha Isomerization Unit & New FCC Gasoline
Selective Desulphurization Unit & Capacity revamp of DHDT
Case-III A: Installation of Units as in Case-I with installation of additional Hydrogen
Unit
Case-III B: Installation of Units as in Case-I with Naphtha import to saturate CCRU
capacity
Crude
Naphtha
Reformate
Euro-IV MS
Euro-VI MS
Base Case
Case-I
Case-II
Case-III A
Case-III B
15000
215
1235
2076
0
15000
230
1408
601
1292
15000
0
1267
0
2263
15000
0
96
0
3260
15000
0
943
0
3881
Apparently, Case-III A seems to be the most suitable option which calls for installation of
new Naphtha Isomerization Unit, new FCC Gasoline Desulphurization Unit and a new
Hydrogen Unit which will be commissioned by 2020. Apart from upgrading the fuel quality,
the proposed scheme envisages maximization of MS from the refinery by absorbing the
entire Reformate produced within the Refinery into the MS Pool.
With the installation of new Hydrogen Unit, the requirement to operate CCR at higher
capacity has been resolved, thereby reducing the Reformate generation. With the
installation of Isomerization and Desulphurization units, the Sulphur is reduced to the
desired level and also reduction in Benzene content of the final MS pool is achieved, thereby
enabling sufficient leverage for CCR unit operation at matching capacity for maximization of
BS-VI quality MS from refinery. Detailed below is the MS pool for BS-VI scenario.
Alkylate
Reformate
Isomerate
Prime-G top cut
Prime-G bottom cut
BLEND
MS Euro-VI
Manufacture SPEC
Mass,
TMTPA
RON
% OLF vol
% AROM
vol
% Benzene
vol
Sulfur,
PPMW
338
1290
846
341
445
3260
96
104
83
89
92
93.8
0.3
0.85
0.1
54
19
8.6
1
82
0.1
3
39
34.3
0.2
1
0
0.9
0.9
0.6
8
0.5
0.5
10
10
3.6
91.6 /
91.2
21%max
35% max
0.9% max
10
4.0 Additional source for MS production
As discussed earlier, PRU LPG is highly olefinic and the total flow to the Alkylation unit is
restricted by the olefin content of the stream and hence, this stream is partly routed to
Alkylation unit and partly to LPG pool. This olefinic LPG consists of iso-butene which by
indigenous technology (Octamax Unit) produces iso-octene having very high octane
number. Following two cases have been studied for production of iso-octene from LPG.


Octamax unit in parallel to Alkylation unit – In this configuration, the component of
C4 stream from PRU, which is going to LPG pool, is routed to Octamax unit and the
balance quantity of C4 stream will be feed to Alkylation unit. The feed available to
Octamax will be limited in order to maintain the RVP of the final LPG pool within the
manufacturing specifications of 1050 kPa (g) with marginal increment of MS
production.
Octamax Unit in series to Alkylation unit – In this configuration, the entire olefinic
C4 stream will be routed to Octamax unit and the effluent stream from Octamax unit
(rich in olefinic C4) will then be routed to Alkylation Unit. This configuration enables
utilization of full potential of the olefinic C4 for production of MS blend component
from Octamax and reduction in olefins increases the utilization of Alkylation unit.
Opportunity exists to further increase MS production from the refinery by separating the
iso-butene from PRU C4 stream going to Alkylation Unit as feed, for production of isooctene, which can form a high RON MS blend component from LPG. This reduction in olefin
content in Alkylation feed has a further advantage of pushing more saturated feed into
Alkylation Unit thereby achieving capacity utilization of the unit and maximizing production
of Alkylate.
In Octamax unit, iso-butylene dimerization takes place to produce iso-octene and the
remaining olefinic and saturated C4 components exit the unit unreacted from the unit which
is then fed to Alkylation unit. This change in routing of the olefinic LPG has dual advantages.
Firstly, this iso-octene produced forms an excellent MS blend component because of its high
RON, zero Aromatics, Benzene & Sulphur. Also, since there exists cushion in the olefin
content of the final MS pool after BS-VI Scenario, the iso-octene will increase the total MS
quantity.
Secondly, this reduction in olefin content of the PRU LPG increases the operating capacity of
Alkylation unit by increasing the overall saturates to the unit, thereby increasing the Alkylate
production and overall MS quantity.
Impact of Octamax
4000
3700
3500
3260
3000
2500
2076
2000
1500
896 932
1000
500
645
377 338 487
0
0
198
610
0 110
0
ALKYLATE
OCTAMAX PDT
Base Case (Fuel Refinery + PP)
MS
Post BS-VI Project
LPG
DELTA GRM
Post BS-VI with Octamax in series before Alky unit
* Quantity in kTPA and GRM in Crores
The revised MS blend in this scenario is as follows:
Alkylate
Reformate
Isomerate
Prime-G top cut
Prime-G bottom cut
Iso-Octene
BLEND
MS Euro-VI
Manufacture SPEC
Mass,
TMTPA
RON
% OLF vol
% AROM
vol
% Benzene
vol
Sulfur,
PPMW
487
1386
846
341
445
198
3700
96
104
83
89
92
114
95.6
0.3
0.85
0.1
54
19
100
12.9
1
82
0.1
3
39
0
32.7
0.2
1
0
0.9
0.9
0
0.6
8
0.5
0.5
10
10
0
3.5
91.6 /
91.2
21%max
35% max
0.9% max
10
5.0 Conclusion
1) With the implementation of BS-VI Projects, besides fuel quality upgradation, there is
an increase in MS production from 2076 to 3260 TMTPA
2) With the addition of Octamax, there will be further increase in MS production to
3700 TMTPA.
3) The addition of Iso-octene into the MS pool produces Premium quality MS from the
refinery.
********************************
H2 recovery fro
MP Separators’ vapor strea s of DHDT u it
Soumitro Majumdar, Malay Bagchi, K K Mandal
Indian Oil Corporation Limited, Panipat Refinery, Panipat, Haryana, India
Abstract
In Panipat Refinery DHDT unit, H2 rich sour vapor streams from Hot MP Separator & Cold
MP Separator are sent to the Stripper overhead line under pressure control of individual
separator vessels. The mixed vapors are partially condensed first in the Stripper Air
Condenser, then in the Stripper Trim Condenser and finally separated in the Stripper Reflux
Drum. The sour vapor from the Stripper Reflux Drum is mixed with the off gases from
Stabilizer & other refinery units and the combined sour gas is routed to the LP Amine
Absorber (6.6 kg/cm2g). The amine treated H2 rich gas is finally routed to the fuel gas
system resulting in considerable loss of H2.
I order to re o er this H , pro ess s he e o H re o er fro MP Separators apor
strea s of DHDT u it as e e uted a d o
issio ed i Fe 6. The H ri h sour apor
streams from Hot MP Separator & Cold MP Separator are mixed under pressure control,
cooled in MP Separators Vapor Water Cooler and sent to MP Separators Vapor KOD. The
condensate is sent under level control to Stripper Reflux Drum and the off gas is sent
directly to OHCU Amine Treater (23.7 kg/cm2g) and then routed to impure H2 header.
The et a ual sa i gs
₹25.97 crores per annum.
re o eri g H fro
MP Separators
apor strea s is around
Original System
In Panipat Refinery DHDT unit, H2 rich sour vapor streams from Hot MP Separator & Cold
MP Separator are sent to the Stripper overhead line under pressure control of individual
separator vessels. The mixed vapors are partially condensed first in the Stripper Air
Condenser, then in the Stripper Trim Condenser and finally separated in the Stripper Reflux
Drum. The sour vapor from the Stripper Reflux Drum is mixed with the off gases from
Stabilizer & other refinery units and the combined sour gas is routed to the LP Amine
Absorber. The treated off gas is finally routed to the fuel gas system at low pressure. As per
original unit design, a two-stage Off Gas Compressor was envisaged for increasing the
pressure of the treated off gas for routing it to HGU for H2 recovery. However, the Off Gas
Compressor Section was not installed during unit commissioning and was proposed to be
installed at a later date. Due to the absence of the Off Gas Compressor Section in DHDT unit,
the amine treated H2 rich gas was routed to the fuel gas system resulting in considerable
loss of H2.
Page 1 of 5
TO STRIPPER OVHD.
PV
PV
HOT MP
SEPARATOR
180 °C
20.0 kg/cm2g
COLD MP
SEPARATOR
60 °C
22.0 kg/cm2g
LV
LV
FV
TO STRIPPER FEED
1ST PREHEATER
FV
TO STRIPPER FEED
2ND PREHEATER
STRIPPER
FEED/BOTTOM
PLATE HE
175 °C
19.5 kg/cm2g
FEED SIDE
DESIGN
PRESSURE
24.2 kg/cm2g
Process Study
As per original design basis, the potential of recoverable H2 from the combined vapor
streams of Hot MP Separator & Cold MP Separator was around 580 kg/hr (6.5 kNm3/hr).
The combined vapor stream had a significant H2 potential and was available at a very high
H2 purity of 80%. In order to recover H2 from these vapor streams, an in-house simulation
was done by adopting the following methodology:
 The H2 rich sour vapor streams from Hot MP Separator & Cold MP Separator shall be
routed for amine treatment in OHCU and finally routed to impure H2 header instead of
routing the streams to DHDT Stripper overhead.
 The existing LP gas treating unit in OHCU receives the combined LP gases from OHCU &
HCU having a H2 content of 81-82 vol%. The amine treated gas having a H2 content of
88-89 vol% is then routed to HGU PSA under fixed flow control and balance gas to
impure H2 header. The impure H2 is then consumed in refinery hydroprocessing units.
The LP gas treating unit was operated at only 70% of its design capacity, thereby
providing a margin for processing the off gas from DHDT.
 As inlet pressure of the LP gas treating unit in OHCU is 25.0 kg/cm2g, the operating
pressures of Hot MP Separator & Cold MP Separator shall be increased to 28.7 kg/cm2g
from the current operating pressures of 20.0 kg/cm2g & 22.0 kg/cm2g respectively. This
required the replacement of the original vessels with new ones as the MAWP of the
original Hot MP Separator & Cold MP Separator is 22.0 kg/cm2g & 24.2 kg/cm2g
respectively.
Page 2 of 5
 An increase in the operating pressures of Hot MP Separator & Cold MP Separator shall
result in an increased pressure on the feed side of the Stripper Feed/Bottom Plate Heat
Exchanger beyond the design pressure of 24.2 kg/cm2g. This required a differential
pressure control valve on the liquid from the Hot HP Separator and protection measures
(PSVs) to avoid any change in the operating parameters of the Plate Heat Exchanger.
 An additional cooling cum separation system for the combined vapor streams. This
required a new water cooler & a new K.O. drum plus associated piping and
instrumentation.
 An increase in the operating pressures of Hot MP Separator & Cold MP Separator also
required the replacement of the associated piping and instrumentation (control valves,
PSVs etc.).
The scheme was shared with Licensor and after receiving feasibility confirmation, Licensor
was awarded the job of process study.
Modified System
The scheme was executed at a total cost of Rs 4.56 crores. The execution of the scheme
involved the following major modifications:
 New Hot MP Separator & Cold MP Separator as the new operating pressure of the MP
Separators is exceeding the design pressure of the original vessels.
 New MP Separators Vapor Water Cooler for cooling the combined vapor stream from
MP Separators.
 New MP Separators Vapor K.O. Drum to separate the desired vapor phase from the
mixed phase cooler effluent and send the liquid phase to the Stripper Reflux Drum under
level control of the new K.O. Drum.
 New piping and instrumentation (control valves, PSVs etc.) in view of increased
operating pressure of MP Separators and additional equipment.
Page 3 of 5
TO MP SEPARATORS
VAPOR WATER COOLER
PV
TO STRIPPER OVHD.
PV
PV
PV
HOT MP
SEPARATOR
180 °C
28.7 kg/cm2g
COLD MP
SEPARATOR
60 °C
28.7 kg/cm2g
LV
LV
FV
TO STRIPPER FEED
1ST PREHEATER
FV
TO STRIPPER FEED
2ND PREHEATER
STRIPPER
FEED/BOTTOM
PLATE HE
175 °C
19.5 kg/cm2g
PDV
FEED SIDE
DESIGN
PRESSURE
24.2 kg/cm2g
The scheme was successfully o
issio ed i Fe 6. The scheme has eliminated the
requirement of an Off Gas Compressor Section in Panipat Refinery DHDT unit as
contemplated in original unit design.
As per study design basis, the H2 rich off gas from DHDT shall have a flow rate of around
1700 kg/hr with a H2 content of 86 vol%. The potential of recoverable H2 from this off gas is
around 520 kg/hr (5.8 kNm3/hr). In actual operation, around 1000 kg/hr of DHDT off gas
having H2 purity of around 95 vol% has been routed to OHCU due to capacity limitation in
LP gas treating unit and present impure H2 consumption scenario. Thus, around 545 kg/hr
(6.1 kNm3/hr) H2 has been recovered through commissioning of this scheme. Additional H2
recovery can be realized if balance DHDT off gas can be routed to OHCU (complete stoppage
of DHDT MP Separators vapor routing to DHDT Stripper overhead) after capacity
augmentation of OHCU Amine Treater. Capacity augmentation of OHCU Amine Treater
through tray replacement along with capacity augmentation of HGU PSA through adsorbent
replacement has already been initiated by Panipat Refinery in order to recover H2 from the
balance DHDT off gas.
Economic Benefit Assessment
The net annual savings by recovering around 545 kg/hr (6.1 kNm3/hr) H2 works out to be
₹25.97 crores with an IRR of 96.24%. The net annual savings has been computed based on
Page 4 of 5
cost of reduced naphtha consumption in HGU minus cost of increased IFO consumption on
account of lower calorific value of FG.
Conclusion
The scheme has helped in recovering around 545 kg/hr (6.1 kNm3/hr) H2 from DHDT unit
through most economically viable option resulting in net annual savings of ₹25.97 crores.
Page 5 of 5
Authors: Anuj Seth and Claus Brostrom Nielsen, Haldor Topsoe
Title: Squeeze more value out of your FCC gasoline
Abstract
With the Ministry of Petroleum and Natural Gas (MoPNG) announcing a nationwide supply of BS
VI fuel (Gasoline & Diesel) in conjunction with the proposed BS VI emission standard
implementation date of 1 Apr 2020, the Indian sulfur legislations tighten. This new sulfur legislation
pose challenge for Indian refiners to meet the specifications from their existing FCC gasoline posttreatment units while utilizing conventional FCC post-treat catalyst. As refiners gear up to upgrade
their gasoline production units for the new legislation, they face multiple challenges, like capacity
augmentation with new product specifications, stream blending in gasoline pool, and aligning the
turnaround schedules for the FCC naphtha post-treatment units with overall refinery shutdowns.
With the increase of FCC naphtha post-treatment units being build, several refineries now face the
challenges of keeping the octane number in the gasoline pool. When hydroprocessing cracked
naphtha feed, olefins to some extent also get hydrogenated resulting in to loss of octane number.
This will prevent the refineries from achieving maximum value of their product. One of the most
efficient way to overcome this challenge is by use of catalyst with optimized activity and
selectivity which can enable removal of sulfur to ultra-low levels while retaining high octane
numbers.
This paper covers Topsoe’s series of HyOctane™ catalysts which are specifically developed for all
steps in FCC gasoline post-treatment applications. Their optimized activity and selectivity enable
removal of sulfur to ultra-low levels while retaining high octane numbers, allowing a stable,
flexible, and profitable production of high-quality gasoline. The HyOctane™ catalysts build on
Topsoe’s advanced preparation technologies and exhibit stable performance throughout the
entire catalyst cycle. The high-activity catalysts enable operation with extended cycle lengths
and/or increased throughput, and high selectivity ensures minimum octane loss. In HyOctane™
catalysts utilizing the in-depth knowledge about the reaction kinetics, selectivity towards sulfur
compound
conversion
has
been
significantly increased while
minimizing
monoolefin hydrogenation. This ensures that a high octane number is maintained in the final gasoline
product. Currently, Topsoe has three catalysts which are commercially available in the
HyOctaneTM portfolio: TK-703 HyOctaneTM (for selective hydrogenation) and TK-710 HyOctaneTM
& TK-747 HyOctaneTM (for first & second stage HDS respectively).
Introduction
The maximum sulfur contribution in the refinery motor spirit (MS)/gasoline pool is the MS coming
from the FCC unit which is basically the
cracked naphtha. It, therefore, depends upon
the sulfur levels in this cracked naphtha, that
what shall be the final sulfur content in the
refinery finished MS product (refer figure 1).
Refineries around the world have two ways to
control the sulfur content in the cracked
naphtha from FCC (1) Either install an
Vaccum Gas Oil (VGO) hydrotreater
upstream of the FCC unit to remove sulfur
from the VGO feeds, or (2) Install a FCC
post-treatment unit for the cracked naphtha.
In Indian context, it is seen that route (2) has
predominated the sulfur reduction strategy
for the MS pool. Now let us compare the two
routes above in little detail.
Fig.1
Figure 2 shows the relationship between the sulfur in the FCC feed and the sulfur in the gasoline
product from a FCC unit. With a low gasoline sulfur specification target of 500 ppm, the feed to the
FCC is in the order of 1 wt%. However, the pre-treater (VGO treater upstream of FCC) severity
needs to be increased dramatically in order to meet the very low sulfur specifications that a refiner
shall target. For FCC gasoline sulfur specification of around 50 ppm, the VGO sulfur content in the
feed to the FCC, needs to be reduced to an order of 500 to 1000 ppm. Such reduction in FCC
feed sulfur are possible in the pre-treater by either increasing the catalyst volumes in the VGO
treater or by increasing its operating severity specifically with respect to temperature.
Fig.2
It was discussed earlier that FCC naphtha represents up to 30-50% of the overall gasoline pool in
a typical FCC refinery. The FCC naphtha should thus contain less than 80-100 ppmw sulfur to
meet a specification of 50 ppmw sulfur in the total pool. For further reduction in sulfur levels the
severity of operation for pre-treater shall have to be increased. The actual degree of required
desulfurisation will also depend on the cut point of the FCC gasoline since most of the sulfur
species are found in the heavy-end boiling range of the FCC gasoline. Lowering the FCC gasoline
cut point can however, considerably reduce the gasoline yield.
With refineries targeting a MS sulfur specification of 10 ppmw, there shall be a strong effect on the
operation philosophy of the current FCC and the pretreating units. The FCC feedstock must be
hydrotreated to a sulfur level below 300-500 wt ppm to comply with the 10 wt ppm sulfur level
without any investment in gasoline post-treatment technologies. The key is to maintain a high
degree of nitrogen removal whilst extending the degree of desulfurisation. Selection of the best
catalyst for a given service will depend on operating pressure, sulfur compounds present in the
feed, and the nature and amounts of inhibitors to the desulfurisation reaction.
Molecular understanding of FCC naphtha
It is important to understand the molecular composition of FCC cracked naphtha before we talk
about the desulfurization of the cracked naphtha. If we look at the boiling points of the sulfur
species like mercaptans, thiophenes and their derivatives, we can classify the naphtha into two
broad ranges, light cracked naphtha (LCN) and heavy cracked naphtha (HCN). For the
differentiation, it can be said that the fraction boiling above 80-85 °C, can be constituted as HCN.
Further it can be seen that there exists a strong relation between the cumulative sulfur level and
the corresponding fractions of cracked naphtha, refer figure 3. Figure 4 presents further details on
carbon numbers present in the cracked naphtha from a FCC unit.
Fig.3
Fig.4
The above understanding is quite important to go into designing strategy of the post-treatment
process for cracked naphtha. Since a majority of the above molecules are olefins (di-olefins &
mono olefins), it is a challenge to remove the sulfur molecules while avoiding the saturation of the
olefins which are the major contributor to the octane number of the motor sprit. This can be
understood by the schematic as in figure 5.
Fig.5
Post-treatment of FCC naphtha
With the fundamental understanding of the molecular composition of the cracked naphtha feeds, it
is imperative that post-treatment of FCC naphtha shall involve a two-step process, selective
hydrogenation (SH) of di-olefins and then a hydrodesulfurization (HDS) step to convert the sulfur
species. The first step of selective hydrogenation is carried out at low temperature compared to
the HDS step and therefore the selective hydrogenation also mitigates the risk of fouling in the
downstream HDS reactors. The selective hydrogenation of the cracked feed also transforms the
light sulfur species into heavy sulfur. A usual layout of the post treatment unit consists of a
selective hydrogenation unit (SHU), a splitter to separate the LCN & HCN and then a hydrodesulfurization (HDS) unit. There may be a mercaptan control unit to have final finishing step. A
schematic is shown in figure 6 below.
Fig.6
Pilot plant test with Topsoe’s HyOctaneTM series of catalyst
TK-703 HyOctaneTM for selective hydrogenation unit (SHU)
Several pilot plant tests were conducted with the Topsoe’s FCC naphtha post-treatment catalyst
TK-703 HyOctaneTM and a detailed analysis was done for the feed while analyzing the component
molecules using a gas chromatograph. The feed analysis chromatograph is shown in figure 7.
Further the product from the SHU reactor was analyzed for the product sulfur species. The results
were quite convincing to prove that TK-703 HyOctaneTM could efficiently convert the lighter
mercaptans into heavier mercaptan molecules. Therefore, it can be seen that with TK-703
HyOctaneTM, a part of LCN sulfur can be converted into HCN Sulfur, thus preserving the olefins in
LCN, which can be then separated in a splitter. Post this step, HCN can be hydrodesulfurized and
the octane of the LCN stream can therefore be maintained at a higher value. The product species
analysis is shown in figures 8 & figure 9. The LCN fraction after treatment with TK-703
HyOctaneTM is practically sulfur free.
Fig.7
Fig.8
Fig.9
TK-710 HyOctaneTM & TK-747 HyOctaneTM for Hydro desulfurization
Topsoe has conducted several pilot plant tests with the Topsoe’s latest FCC naphtha posttreatment catalyst TK-710 HyOctaneTM and TK-747 HyOctaneTM utilizing a HCN feed and while
maintaining the test conditions, as in an industrial FCC Gasoline post-treatment unit. For such
tests the cracked naphtha feed was first selectively hydrogenated in the SHU reactor (R1) and
then split into LCN & HCN feeds. The HCN feed was then fed to SHU reactors in series (R2 &
R3).
Topsoe ran a number of pilot tests while testing two combinations: (1) Employing only one HDS
reactor using CoMo catalyst TK-710 HyOctaneTM and, (2) Employing two HDS reactors using both
CoMo catalyst TK-710 HyOctaneTM and the Ni catalyst TK-747 HyOctaneTM. It can be seen from
the two curves (Refer figure 9) that a higher sulfur conversion can be obtained with the two HDS
reactor system while having the same olefin conversion.
Going a step further, for a commercial unit, the performance data from a test run was collected
and analyzed together with the above pilot plant studies. The results were represented in the form
of a plot. The results in form of (black circled square in figure 9) from the commercial test run, can
be used to infer that test run results match with the results from the pilot plant test, since the
results fall into the same curve which demonstrates the proven ness our HyOctaneTM catalyst.
Fig.10
Conclusion
Topsoe’s series of HyOctane™ catalysts are specifically developed for all steps in FCC gasoline
post-treatment applications. Their optimized activity and selectivity enable removal of sulfur to
ultra-low levels while retaining high octane numbers, allowing a stable, flexible, and profitable
production of high-quality gasoline. Currently Topsoe has three catalysts which are commercially
available in the HyOctaneTM portfolio: TK-703 HyOctaneTM (for selective hydrogenation) and TK710 HyOctaneTM & TK-747 HyOctaneTM (for first & second stage HDS respectively)
The innovative recipe of the HyOctaneTM catalysts has led to high stability, which is a result
achieved through years of research. It has been industrially proven that units using HyOctane™
catalysts experience a significantly improved HDS activity. The refiners can make full use of the
advantages like superior activity for sulfur removal, stable operation allowing for longer operating
cycles, low octane loss and more than 99.9 wt% naphtha yield while utilizing the Topsoe
HyOctaneTM catalyst in their FCC post-treat naphtha units. The HyOctane™ catalysts series is
now available to meet the growing demands for cleaner fuels. In conclusion, the advantages of the
new HyOctane™ catalysts series can be utilized to improve the overall profitability and economy
of all kind of FCC naptha post treating unit in multiple ways. Topsoe is committed to Indian refiners
to offer the best catalyst to fulfil their requirements of meeting 100% BS IV gasoline specification
and 100% BS VI gasoline specification by Q1 2020 in the most optimal way while improving their
refinery profitability.
WHITE PAPER
Model-Based Utilities
Optimization and
Management
Authors: Terumi Okano, AspenTech
Sunil Patil, AspenTech
Jack Zhang, AspenTech
Table of Contents
1.
Introduction ....................................................................................................... 3
Table 1: Flexibilities and Constraints Inherent in Many Utilities Systems .....4
2.
Business Processes........................................................................................... 6
2.1
Demand Forecasting .........................................................................................6
2.2 Utilities Production Planning ..........................................................................6
2.3 Optimal Plant Operation (Online Optimization) .......................................6
2.4 Performance Monitoring (Utilities Equipment) .........................................7
2.5 Investment Planning .........................................................................................7
2.6 Emissions Monitoring, Management, and Trading ...................................7
2.7 Contract Management .....................................................................................8
2.8 Tariff Evaluation..................................................................................................8
2.9 Cost Accounting ................................................................................................8
2.10 Power and Fuel Trading ....................................................................................8
2
3.
Technology ........................................................................................................ 9
4.
Example Applications: “Energy Management in Action” .......................... 10
5.
Summary........................................................................................................... 11
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
1.
Introduction
With the ongoing uncertainty of oil prices, rig counts are down, exploration is down and new
pipelines are generally not being built. In gas plants, producers need to find ways to increase yield
and continuously meet stricter quality and environmental regulations. Refineries operate in a market
with fluctuating demands and multiple feedstock options. Refineries also typically operate on tight
profit margins, making margin analysis crucial. As a direct result, in all these areas there is a need
to improve margins and increase asset utilization. Business focus in this current climate is on cost
performance leadership, operational risk reduction, and maximum return on capital employed. To
this end, improving energy efficiency becomes an important part of the equation and it is necessary
to be proactive to the changes in the business environment.
Energy costs are significant for refineries; for a typical refinery, energy is estimated to cost between $75M
and $140M USD per annum. When considering energy efficiency improvements of a processing site,
companies tend to focus pr imordially on process improvements and modifications to the process units
to reduce the amount of energy used by the site. This is known as demand-side energy management.
However, making improvements to the way energy is sourced and supplied can also reduce the amount of
energy used at a site. This is supply-side energy management or utilities management.1
In most sites the operation of the utilities system and its costs are considered a necessary evil:
something that is required to operate the processes, but not given the same level of attention as the
process units. Utilities are not typically a profit center as utilities costs are spread proportionally over
the different process units, resulting in hidden energy inefficiencies. An Integrated Utilities System
Simulation and Optimization environment, specially designed to address all the business processes
related to the operation and management of industrial energy and utility systems, brings all the
pieces together (Figure 1).
Figure 1: An Integrated Utilities System Simulation and Optimization Tool Drives Optimal Decision-Making
1
Supply-Side vs. Demand-Side Energy Management. Fortunately, savings in demand-side and supply-side are usually additive; however,
insufficient consideration on the supply-side can reduce or even nullify the impact of improvements in energy efficiency on the
demand-side. For example, steam savings in one process unit can lead to venting elsewhere if the steam system is not balanced.
3
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
Model-based decisions can be made at the strategic, tactical, and operational levels. Not having
a model which captures the many complexities of utilities systems in a single view means there
is significant untapped potential for improving the energy efficiency of a site through utilities
management and optimization.
Figure 2: Example Utilities System
Table 1: Flexibilities and Constraints Inherent in Many
Utilities Systems
Utilities System Flexibility / Constraints Opportunities
for Optimization
4
Optimization Positively Impacts:
Efficiency
Emissions
Costs
Choice of boiler; supplementary firing of heat recovery
steam generator
x
x
x
Use of auxiliary equipment, like BFW pre-heating
x
x
x
Electricity & fuel trading
-
x
x
Operation of multi-extraction turbines
x
-
x
Steam balance: mismatch between steam production and
consumption at the various levels
x
x
x
Fuel balance: e.g. site may be generating too much
fuel gas for optimal steam production, resulting in the
inefficient sending of fuel gas to flare
x
x
x
Choice of drives in process units (turbine/motor) creating
discrete steps in steam demand/consumption on specific
levels
x
x
x
Variability of utilities demand/production from process
units; ability to forecast process demand
x
x
x
Emission constraints on greenhouse gas and carbon
dioxide trading
x
x
x
Equipment availability/capacity constraints/maintenance
-
x
x
Utilities contractual constraints: tiers/nominations/
penalties, etc.
-
-
x
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
Figure 2 shows the main components of the fuel and steam networks of a simple utilities system,
including the boilers, gas turbine, heat recovery steam generator, turbo-generators, pump drives,
letdown stations, and deaerators. Table 1 shows the flexibilities and constraints such a system
generally exhibits and the effect of optimization on energy efficiency, emissions, and energy costs.
Utilities systems in the process industry vary considerably in size and complexity and, as outlined
in the previous paragraphs, every utilities system has a site-specific combination of flexibilities and
constraints. Therefore, the best approach to achieve optimum operation of the utilities system can
also vary significantly.
As a starting point for achieving optimum energy efficiency, several guidelines can be followed,
such as minimum static letdowns, no steam venting to atmosphere at any level, and maximum use
of extraction turbines. These simple guidelines might be sufficient for small, non-complex sites, or
sites with limited operational flexibility. With increasing complexity, however, optimum and energyefficient operations can only be achieved by using the right tools. These can range from spreadsheetbased simulation tools or distributed control system (DCS) programming to more powerful modelbased offline or online optimization tools, and might be integrated with other manufacturing and
execution systems on-site.
The following sections describe how a model-based utilities management and optimization system
(a Utilities Optimizer) can successfully be deployed in complex sites through multiple business
processes.
5
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
2. Business Processes
The challenge of operating an energy-efficient utilities system can best be addressed by breaking it
into a series of business processes, each of which should be supported by consistent data and advice
provided by a single system. Some of the business processes are largely driven by energy efficiency,
while others are predominantly concerned with energy costs. The following subsections describe
various business processes that can be supported by a model-based Utilities Optimizer. In addition,
site-specific business processes like load-shedding or fuel management can be supported. The mix
of business processes implemented depends on the benefits predicted for each.
A selection of the more commonly implemented business processes is described in more detail
below.
While the Utilities Optimizer will improve the site’s energy efficiency, it may help even more in
reducing the site’s energy bill. Cost reduction is a strong driver for installing this solution at an
industrial site. With a Utilities Optimizer, attention has been given not only to energy efficiency, but
also to the other key factors – such as energy costs and emissions – that should be considered when
determining the economic feasibility of installing the solution at a site.
2.1
Demand Forecasting
To operate and manage the utilities supply system at the lowest cost requires knowledge of the
current and predicted future utility demands. This helps to minimize the use of hot standby (e.g.,
boilers), the venting of steam due to excess online capacity, and the loss of supply due to insufficient
standby or control. It also ensures that penalties are not incurred due to violation of take-or-pay
contracts, maximum demand charges, or load factor clauses in both the electricity and gas contracts.
2.2 Utilities Production Planning
Utilities production planning involves taking the demand profiles and, based on availability of utilities
generation equipment, developing an optimized production plan within the constraints of the utilities
tariffs. Production planning can typically be carried out on both a tactical and strategic basis. The
tactical plan would be concerned with the next 24 hours, while the strategic plan, for example, might
be concerned with the best configuration for when a gas turbine needs to be brought down for
inspection or when a process unit is to be shut down or started up.
2.3 Optimal Plant Operation (Online Optimization)
While a plan may be developed in advance, in practice the operations of the plant may change within
that period, thereby invalidating the optimum plan. Even if the steam and power demands do not
change within this period, other factors such as electricity price and gas price may vary. The Utilities
Optimizer can provide real-time advice to operations personnel on how to best configure and
operate the system at the lowest cost based on current demands and prices. A side benefit is that
the operator is also able to use a real-time system to perform what-if analysis to evaluate alternative
operating modes, for example to cater to a unit shutdown.
6
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©2017 Aspen Technology Inc. AT-03163
2.4 Performance Monitoring (Utilities Equipment)
If the necessary metering is in place, the Utilities Optimizer can track performance of the individual
items of equipment within the utilities system, such as the efficiency of the boilers and gas turbines.
This information can be used to optimize cleaning and maintenance schedules and can also provide
early warning of operating problems.
2.5 Investment Planning
The model-based Utilities Optimizer can be used to evaluate design options – for new equipment
and changes to existing equipment – in both the process units and utilities systems to improve the
overall energy efficiency of the site. Examples for these types of improvements include:
• Using process heat for heat deaerator feed water
• Choosing drives (motor or steam turbine) or dual-process drives for greater flexibility in balancing
the steam system
• Changing energy supply (e.g., using low pressure steam to reduce medium pressure steam use)
For analysis of these opportunities, the Utilities Optimizer can be used to support the activities of a
site-wide energy strategy study, incorporating a series of standard models. The engineer can use this
to accurately evaluate the plant-wide economic benefits of any change to the current configuration,
within the process units or utilities system. This provides the project engineer with the confidence
that the best information and current system ability is used when evaluating investment options. In
turn, this ensures that the limited capital is wisely spent and minimizes regret capital.
2.6 Emissions Monitoring, Management, and Trading
Increasingly, a permit to operate requires that plants operate within strictly defined environmental
constraints. When the permit to operate specifies a maximum emissions limit, this causes a
constraint to production, the sourcing of energy, and the operation of the utility system. In such
cases, it is vital that a Utilities Optimizer includes emissions prediction and reporting.
Depending on (local) legislation, emissions management and trading is a necessary and important
business process. The cost of emissions can influence the choice of fuels and the make vs. buy
balance for satisfying the electricity demand of the site. A Utilities Optimizer predicts2 the likely
energy demands and corresponding emissions across a range of potential operating scenarios,
supporting decision-making for trading of emission allowances.
7
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
2.7 Contract Management
Once a tariff has been selected and a contract has been established, there is an ongoing requirement
to manage the utilities consumption within the terms of the contract.
Often the key to driving down the cost of purchasing utilities is reducing both the average and
maximum demand, in addition to exploiting the tariff structures to the advantage of the operating
company. In general, the bigger the gap between the average and maximum demand, the larger
the unit cost of utility (electricity or gas). Companies can exploit degrees of freedom within the
site operations to stay within the contract constraints. For example, flexibility can be exhibited
by switching off electric motors and switching on steam turbines or burning fuel oil rather than
imported natural gas. The Utilities Optimizer can provide the operator with an accurate picture of
the current operation as well as likely future demands and operation, thereby highlighting costly
potential problems. The operator is then able to use the system to identify the best way to prevent
the problem from occurring.
2
This is based on demand forecasting.
2.8 Tariff Evaluation
In recent years, utilities markets around the world have become deregulated and opened up to
competition. One natural consequence of this is that the process industry is finding new options
to reduce utilities costs by contracting with different suppliers and through new types of supply
contracts. As such, the site operator is now faced with a bewildering array of tariff options. Making
the right choice of supplier is not simply a case of choosing the lowest unit cost of utility (e.g.,
cents per kWh). Most tariff structures include elements of maximum demand charges and punitive
penalties for exceeding maximum demand.
Without the benefit of suitable software, the task of tariff selection can be very labor intensive
without any guarantee that the best option will be selected. Companies operating large cogeneration
systems (which are increasingly the norm) have the added complexity of being able to export
electricity.
2.9 Cost Accounting
In many companies the allocation of costs for utilities can be somewhat arbitrary and hence
unreliable.
A Utilities Optimizer provides the ability to perform accurate cost allocation, including the provision
of utility costs in real-time. It also provides true marginal costs, which apply to reduced or increased
use of utilities. Real-time prices can support decisions such as increasing steam use to a particular
unit to enhance throughput or determining the cost of steaming-out of equipment.
2.10 Power and Fuel Trading
While there is no doubt that significant value can be gained by optimal trading in utilities, likewise
sub-optimal trading can increase risks and incur significant cost penalties. The key to optimal trading
lies in knowing the exact current position and having the ability to deviate from this, both in terms of
technical capability and cost. The Utilities Optimizer can provide optimal trading information online
to the power and fuel traders, therefore supporting efficient trading.
8
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
3. Technology
The Utilities Optimizer explores the flexibility inherent in the purchase, generation, use, and
distribution of utilities, advising the user on the optimum choice available. As previously stated, a
Utilities Optimizer enables all of the following factors to be considered in the optimization: different
tariffs; alternative fuels; optimum loading of boilers and turbines; choice of equipment; import, selfsufficiency or export of electricity; and choice of drives (motor or turbine).
A Utilities Optimizer provides a model-centric approach whereby a single rigorous model of the
utilities system is used to address all of the important business processes.
From a key technology requirement standpoint, a model-based Utilities Optimizer should:
• Offer a flexible modeling and optimization capability that accurately represents the fuel, steam, and
power generation processes and distribution system
-
Properties of all fuels, including lower heating value and composition
-
Thermodynamic properties of all water and steam streams on the facility
-
Performance of all utility equipment over their normal range of operation
-
Addition of any site-specific contractual or process constraints
• Incorporate the exact cost of each utility imported or exported, considering tiered pricing
structures, take-or-pay clauses, and peak demand charges
• Enable multi-period utilities production planning and utility equipment on/off decisions, as well as
discontinuities in the contract model and/or utilities process model
• Reconcile the utilities heat and material balance due to random and systematic errors in plant
measurements
• Provide real-time actionable guidance to the plant operator to efficiently optimize utilities
operations, taking into account any contractual, process and environmental constraints
• Enable offline and online what-if studies
• Provide process unit utility demand models integrated with the production planning system so that
utility and process production plans are consistent
9
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
4. Example Applications: “Energy Management in Action”
Figure 3 shows typical energy savings across the energy management lifecycle. Major energy users
such as oil refineries and large chemical complexes are typical of the sites that can benefit from
adopting a model-based approach to energy performance management.
Figure 3: Typical Energy Savings with aspenONE® Energy Efficiency Solutions
For example, one of the largest oil complexes in the world, Kuwait National Petroleum Company
(KNPC), modeled both the demand and supply side with Aspen HYSYS® Petroleum Refining, Aspen
Exchanger Design and Rating (EDR), and Aspen Utilities Planner™ to reduce energy use. They
projected a possible savings of over $15M USD per year. Similarly, Shanghai SECCO Petrochemical
Company used Aspen Utilities Online Optimizer™ to minimize their utilities costs and achieve their
corporate goal of reducing energy consumption and emissions. They did this by optimizing the fuel
mix and boiler selection, conducting steam balancing, and evaluating the use of turbines versus
motor drives.
KNPC also put in place online performance monitoring for CO2, evaluating equipment efficiencies
and process specific energy use and costs. They lowered total operating costs by more than 1
percent, maximized boiler loads and internal power generation during peak times, and minimized
internal power generation during off peak times. Operational change opportunities were discovered
that could be taken advantage of with nearly no costs.
As one final example, due to significant business changes, Mitsui Chemicals developed a complexwide utilities model along with a Microsoft® Excel interface. This enabled operators to execute
the system from the DCS so they didn’t need to run Aspen Utilities Planner themselves. Mitsui
Chemicals achieved a total operational cost reduction of 1.5 percent.
10 Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
5. Summary
In today’s environment of uncertainty, model-based utilities optimization and management can help
process manufacturers make the best strategic, tactical and operational decisions. Energy-intensive
processes combined with the need to be proactive makes understanding energy use on both the
demand and supply side a necessity. A typical Utilities Optimizer3 will enable companies to:
• Reduce energy costs across sites globally by up to 2-5 percent
• Improve the way energy is sourced, traded, and used
• Understand fuel, steam, and power usage across site(s)
• Identify tradeoffs between increasing production vs. increasing energy costs
Additionally, a Utilities Optimizer enables companies to obtain real-time operational guidance with
regards to:
• Changes in feedstock availability or grades
• Price variation
• Environmental regulation
• Changes in market demand and other unexpected operational changes
Business focus in this current climate is on cost performance leadership, operational risk reduction,
and maximum return on capital employed. To this end, improving energy efficiency becomes critical
to succeeding in today’s business environment.
3
11
See Aspen Utilities at http://www.aspentech.com/products/engineering/aspen-utilities-planner/
Predict, Prescribe, Profit: Creating a World that Doesn’t Break Down
©2017 Aspen Technology Inc. AT-03163
AspenTech is a leading supplier of software that optimizes process manufacturing — for energy, chemicals, engineering and
construction, and other industries that manufacture and produce products from a chemical process. With integrated aspenONE®
solutions, process manufacturers can implement best practices for optimizing their engineering, manufacturing, and supply chain
operations. As a result, AspenTech customers are better able to increase capacity, improve margins, reduce costs, and become
more energy efficient. To see how the world’s leading process manufacturers rely on AspenTech to achieve their operational
excellence goals, visit www.aspentech.com.
Worldwide Headquarters
Aspen Technology, Inc.
20 Crosby Drive | Bedford, MA 01730 | United States
phone: +1-781-221-6400 | fax: +1-781-221-6410 | [email protected]
Regional Headquarters
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phone: +1-281-584-1000
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phone: +55-11-3443-6261
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phone: +65-6395-3900
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phone: +973-13606-400
For a complete list of offices, please visit www.aspentech.com/locations
12 Model-Based Utilities Optimization and Management
©2017 Aspen Technology Inc. AT-03163
Application of High Pressure Breech-Lock (Screw Plug) Closure Heat
Exchangers for Refinery Services
Ugrasen Yadav and Amrendra Bet
TECHNIP India Limited
1. Abstract
Shell and tube exchangers are workhorses in the process industry, covering a wide range of
temperatures, pressures and volumetric flowrates, etc., being available in almost any material and
having an established manufacturing technology. Due to robust construction features shell and tube
exchangers can handle high pressure in the range of 90 Kg/cm2 g to 300 Kg/cm2g, but this is not limited
to high pressure but also for high pressure with elevated temperature up to 500°C as well.
High pressure heat exchangers are utilized in the refinery processes like various hydrotreaters, Lube oil
base stock units and for hydrocracking processes. High pressure heat exchangers are expensive to build
as high pressure and elevated temperatures necessitate higher thickness of equipment. Moreover these
severe services often require exotic metallurgy. The selection and design of high pressure heat
exchangers is critical since in the event of maintenance these exchangers could take several weeks.
Fo
efi e
se i es, the te
D-T pe i TEMA is used to des i e specially designed non-bolted
closure for high pressures and high temperatures applications. The term D-type is generalized term
since there are several such designs and some of them are patent. Such closures are costly, however,
and many design engineers prefer to use conventional heads in conjunction with hydraulic bolt
tensioners – the latter being necessary to tighten the very large nuts and bolt associated with high
pressure design. Modified D-Type closure are special closure developed by licensors, commonly known
as Breech lock closure or Screw plug closures. Breech lock / Screw plug exchangers have relatively less
weight compared to conventional D-type heat exchangers. There are two types of these closures, (a)
Internal shell with bundle pull through (differential pressure design), High-High type (HH type) and (b)
Integral channel with tubesheet and bundle, shell pull through – High-Low (HL type). These closures are
often known for its operational flexibility and ease of online maintenance. These special closure ensures
re-tightening of tubesheet gasket during operation thereby minimizing mechanical failures, these
closure ensures avoidance of high stressed zones in pressure parts, relatively small bolts compared to
conventional D-type heat exchangers and all channel parts can be removed without cutting. One of the
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major advantage breech-lock closure (differential pressure design – HH type) is that, it ensures leakage
containment within the exchanger itself thereby preventing fire and safety hazards.
The objective of this paper is to highlight thermal and mechanical design aspects of the High pressure
heat exchangers including the proprietary Breech-lock closure/screw plug closure. This paper has also
included recent advances in proprietary closure design which improves the operational flexibility and
mechanical failures.
2. Introduction
Shell & tube exchangers are most common type of exchangers in refineries. Shell & Tube exchangers are
constructed in accordance with TEMA standard (Tubular Exchanger Manufacturers Association)
requirements and either the requirements of ASME Section VIII Div. I or II. TEMA nomenclature is widely
adopted in the industry for shell & tube exchangers. High p essu e heat e ha ge s use D-t pe , hi h
is also called as special high pressure closure. High pressure heat exchangers are normally designated
with DEU, DFU, DKU, DED etc. Mostly utilizes U-tube design, embodying only one header, one tubesheet
and with lesser number of gasketed joints as compared to straight tube design.
Fig.1 Conventional D-type Heat Exchanger & Breech lock / Screw plug closure
Apart from TEMA, ASME and API many operators or Engineering and Construction companies have their
own specifications which supplements the basic requirements of above codes. In many cases,
supplementary requirements reflect particular project requirements or specific requirements of an
operating plant, e.g. maintenance methods, etc. These are all valid requirements which although
increasing initial exchanger costs -are justified in reduced operating or maintenance cost. In certain
cases, however, requirements are included which reflects problems an individual company encountered
in the past, but with latest technology, this is no longer a problem. To minimize cost, it is necessary for
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specifications to be reviewed on regular basis (preferably in conjunction with fabricators), so that
obsolete requirements can be deleted.
Cost of D-type closure shell & tube heat exchangers are determined largely by the initial specifications,
type of D-type closure, surface area required, material of construction, maintenance considerations and
mechanical considerations. In a shell & tube heat exchanger, the cost of raw materials will result in
between 50-70% of the exchanger selling price, depending on the material of construction selected.
Weight of conventional closures in comparison to high pressure modified D-type (Breech lock / Screw
plug closure) are much higher for same operating and design conditions. Normally weight of the
conventional closure type exchangers is higher in the range of 25-45% depending upon diameter of the
exchanger for given design temperature & pressure.
3. High pressure heat exchangers for Refinery service:
Hydrotreating & hydrocracking plant operates at very high pressure and temperature conditions, also
involving hydrogen rich fluid containing hydrogen sulphide mixtures. Due to high risk of explosion and
corrosive flow conditions, heat exchangers for such services requires highly reliable sealing features.
Also due to corrosive fluid conditions in the exchangers at elevated temperature and pressure
conditions necessitate frequent inspections, which makes it necessary that the exchanger components
are easy to deassemble and reassemble during periodic maintenance & shutdowns.
3.1 Use of special closure exchangers in Hydrotreating units:
The term hydrotreating is used to define treatment processes like, saturating olefins to paraffins (for
product stability), removal of sulfur (hydro-desulfurization), saturation of aromatics (mostly for diesel to
improve cetane number. This is the most severe operation of hydrotreating, to improve color of lube
base oils by removing Nitrogen components. With stricter norms for sulfur for petroleum products,
there are multiple Hydro-desulfurization(HDS) units in any refinery, HDS of light component like naphtha
(to make reformer feed), HDS of diesel, HDS of vacuum gas oil (FCC feed), HDS of residues (like ARDS and
VRDS etc.). Hydrotreating is carried out at high temperature & moderate pressure over two stage
reactors in the hydrogen atmosphere. Depending on feedstock used reactor effluent temperature varies
around ~300°C for Naphtha & around ~390°C for atm residue at moderately high pressure between 90150 kg/cm2g. Pumped feed mixes with make-up and recycle hydrogen (depending upon the specific
licensor scheme) in Reactor feed / effluent exchangers (High pressure heat exchangers). Reactor effluent
is cooled by exchanging heat with reactor feed, distillation feed and an Air cooler.
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Fig.2 Typical Hydrotreating process
3.2 Use of special closure exchangers in Hydrocracker units:
Hydrocracking is process to convert higher boiling hydrocarbon molecules to lower boiling molecules by
simultaneous or sequential hydrogenation and C-C bond breaking. Hydrocracking is carried out at
moderate temperature & high pressure over a fixed bed of catalyst where hydrocarbon feed is cracked
in the hydrogen atmosphere. Feed to multistage pump operating at 65C – 230°C. Discharge pressure can
vary from 100-200 kg/cm2g. Pumped feed mixes with make-up and recycle hydrogen in Reactor feed /
effluent exchangers (High pressure heat exchangers). Reactor effluent is cooled by exchanging heat with
reactor feed, distillation feed and an Air cooler.
Fig.3 Typical Hydrocracking process
There are two types of D-type closures are utilized in process and refinery industry;
I.
Conventional D-type closure
II. Special or modified D-type closure (Breech lock / Screw plug type)
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4. Conventional D-type closure:
Conventional D-type closures need very large thickness flanges, very large size bolts & nuts and very
high strength compression gaskets etc. As there are number of flanges, number of leakage points e.g.
shell flanges, channel and channel & channel cover flanges etc. These heavy flanges and thick/heavy
channels / channel cover leads to very high cost of the conventional D-type exchangers. Use of very
heavy and large size bolts & nuts increases maintenance hassles.
Fig.4 Typical Conventional D-type Heat Exchanger assembly
Disadvantages of conventional D-type exchangers:
I.
Higher sizes of bolts due to higher thickness (due to higher design pressure & temperature) of
channel and channel cover.
II. Conventional D-type closures need lip seal gaskets welding to channel / channel cover flange,
due to this cutting of internal parts such as diaphragm may be required during shut-down
III. Very difficult dismantling / reassembly for inspection during shutdown
IV. Bulky equipment, since both shell side & tube side designed for higher design pressure &
temperature
V. Due to uneven expansion between channel and channel cover, frequent leakage problems
occurs in these conventional type closures
VI. In case of galling of conventional D-type closure, there is no other options to rather cut the bolts
or take out complete. Due to this there will be significant loss of production to the company
since either they need to bypass this exchanger or run the plant at significantly low throughput.
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5. Modified D-type closures (Breech lock / Screw plug closures):
Since the cost of these high pressure heat exchanger is much higher than normal shell & tube heat
exchangers, special closures have been developed by the many licensors, to reduce the cost as well as to
minimize the leakage problems and ease of maintenance of these heat exchangers. These special
closures are called as modified D-type closures or breech lock / screw plug closures. The very first
closure
was
invented
and
developed
in
1960
by
Standard
Oil
Co.
of
California.
The first breech-lock was put into operation in late sixties. Rights of technology remain with Chevron
Research Company of California. In 1982, a USA based process licensing company became a licensee for
the worldwide marketing of high pressure breech lock / screw plug heat exchangers. Since early eighties
few international vendors and companies contributed to the further development of the breech lock /
screw plug internals design. Few international vendors have filed the patents for proprietary
modifications and marketed them further. Now more than 2000 breech lock / screw plug high pressure
heat exchangers are in operation worldwide. There closures find its applications in refinery units
operating at high pressure and temperature with hydrogen rich streams such as hydrotreaters and
Hydrocrackers. For efficient heat transfer, maintenance becomes a very essential requirement and
therefore Breech Lock Heat Exchangers become an attractive option provided they function as intended.
These special Breech lock / Screw plug closure are also of two type, High-High Closure and High-Low
Closure.
5.1 Internal shell with bundle pull through (differential pressure design), High-High type (H-H type),
high tube side pressure-high shell side pressure. It’s a removable tube bundle design, used when
differential pressure is specified in process specifications. Differential pressure design is method which
can be used to reduce the thickness of the tubes, tube sheet and hence save cost.
The following considerations apply for use of differential pressure design: (I) The fluid must always exert
pressure simultaneously e.g. a reactor feed/effluent exchanger where common source for both the
fluids. (II) There must be no valves or restrictions which allow one fluid to exert pressure without other.
(III) Start-up, shutdown and upset conditions must be considered and evaluated. (IV) Sequence of
hydrotesting must be established to ensure both fluids are pressurized simultaneously. (V) In most
cases, a safety device such as bursting disk is used to link the high pressure side to low pressure side. (VI)
Differential pressure design yields substantial cost savings.
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5.1.1 Advantages of H-H closure :
I.
Full containment in case of leakage thereby minimizing fire explosion hazard.
II. The total number of flanged joints are reduced to a minimum due to integral construction of
pressure parts.
III. Integral gasketed joints are adjustable by push rods with bolt ring construction.
Fig.5 Typical Internal shell with bundle pull though High-High (HH) deign closure
Fig.6 Internal shell with bundle pull though High-High (HH) deign exchanger
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5.2 Integral channel with tubesheet and bundle, shell pull through – High-Low (HL type), high tube side
pressure – low shell side pressure. It’s a e o a le shell desig , he e tubesheet is integral part of the
channel assembly. Both shell side and tube side components are designed for the full pressure. A very
little chance of leakage is possible between shell side and tube side fluid in this type of closure.
5.2.1 Advantages of H-L closure :
I.
Channel cover is relatively thin & light
II. Due to diaphragm construction, no cladding of channel cover is required, resulting in cost
savings.
Fig.7 Typical Internal channel with tubesheet and bundle High-Low (HL) deign closure
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Fig.8 Internal channel with tubesheet and bundle High-Low (HL) deign exchanger
5.3 Thermal design features of Breech lock / Screw plug closure:
I.
Longer tube lengths are used for any given surface area for high pressure heat exchangers, so
fewer tubes are needed, requiring less complicated header plate with fewer holes drilled.
Further results in lower shell diameters and lower cost.
II. Ge e al ules of thu
a
e
iolated fo high p essu e heat e ha ge desig such as
bundle weight restrictions, tube lengths & shell ID restrictions etc.
III. No tubes in window (NTIW) design - most commonly used, for vibration free design.
IV. Specially designed supports for the inlet and outlet region for non-NTIW designs.
V. Lower Rho-V-sq values for shell side and tube side in comparison to normal low pressure shell &
tube heat exchangers.
VI. Use ASME Section VIII Div-II instead of Div-I for very high pressure services, results in lower
thickness of components leading to lower weight of equipment.
5.4 Mechanical design features of Breech lock / Screw plug closure:
I.
Reliable high-pressure performance sealing.
II. Easy dismantling / reassembly for inspection.
III. Avoidance of high stressed zones in pressure parts.
IV. Number of flanged joints are reduced to a minimum due to integral construction of pressure
parts.
V. The overall heat exchanger is lighter (thinner channel) and easier to handle.
VI. The internal tube sheet to shell gasket can be tightened from the outside during the operation.
VII. Flexible spiral wound gaskets are used to absorb movements caused by large temperature
differences.
VIII. The hydrostatic pressure load is taken by the channel body and not heavy bolting (bolts are only
sized for gasket compression loading). This allows a relatively thin design of the channel (no
edge bending due to bolting) with consequent reduction in exchangers size and weight.
IX. There is no need for hydraulic bolt tensioning devices to remove the channel cover and
reassemble the exchanger, thus eliminating the number of tightening cycles and strain gauge
measurement.
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X. In cases when internal surface of the channel is overlayed with stainless steel, the risk of
disbanding of weld overlay or other stress induced cracking is minimized due to the elimination
of internal attachment welds, threaded holes in the forgings and tension forces on the weld
overlay.
XI. Smaller bolts compared to simple D-type heat exchangers.
XII. Easy and quick maintenance.
XIII. No threaded holes in forging or in the cladding.
XIV. Internal gasketed joints are adjustable by push rods with bolt ring construction. This results in
relatively small sized bolts and eliminates need for bolt tensioning equipment
XV. No cladding of channel cover is necessary due to the use of the diaphragm cover.
5.5 Advantages of Breech lock / Screw plug closure:
I.
No cutting of welded parts (i.e., channel diaphragm weld) is required to open the channel for
maintenance.
II. Tightening of the gaskets is done by small sized bolts using normal wrenches.
III. No bolt tensioning equipment is required. Therefore, no complicated strain gauge measurement
is required and large number of tightening cycles can be avoided.
IV. Push bolts do not sustain the hydrostatic pressure; consequently, the errors in the tightening
procedures do not endanger the closure integrity. In case of mal-operation only push rods will
get damaged and can be replaced.
V. Time required for disassembly and assembly of closure is reduced due to special construction
features thereby reducing maintenance cost.
VI. Pressure load on the channel cover is absorbed by channel forging through threaded ring
eliminating the need of heavy bolting.
VII. Tube sheet to shell joint can be tightened through external bolting during operation.
VIII. Threaded closures permits smaller diameter of channel as compared to flanged closures
requiring less space when units are stacked.
IX. Hydrotesting of the tube bundle can be performed with channel cover removed.
5.6 Recent developments and recommendations:
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I
e e t ea s’
ajo p o le
e ou te ed
ith the use of
ee h lo k / S e
plug losu es a e (I)
Jamming of the threaded plug, due to deformation of the channel barrel. Thus the opening of the end
closure by unscrewing becomes a difficult task. With the increase in operating temperatures and
pressures, the problems become more severe, due to which, users are not inclined to use these type of
end closures. (II) The leakage through the gasket between the shell and tube-sheet, causing the
intermixing of shell and tube-side fluids. After thorough investigation and analysis, it was found that the
additional forces were acting on the gasket due to thermal expansion of the internals. In order to
resolve first problems, different vendors have followed different approach like modification of channel
design, modification of the internals or modification of material of construction of internals. The later
problem was resolved by vendors by using higher MOC of internals and by providing specific and higher
thickness of gasket.
Some of the manufactures have given up on trying to make these connections operate leak free, and
prefer to weld in the tube sheets and diaphragms. This significantly adds to the cost and time needed to
open and close these exchangers, as special equipment is needed to machine out the parts and weld
them back together again.
When high pressure exchangers are built correctly they provide the highest level of reliability and leak
free service at the lowest total cost of ownership. When they are not built correctly they will leak,
requiring additional plant shutdowns to correct at a pretty substantial cost in lost production. In
addition, there is a pretty significant cost when facilities have to go back to the manufacturer to fix and
correct mistakes made by the manufacturer in the original design, or to make modification to improve
operation or maintainability. Some of the major recommendations were as follows:
I.
Exchangers should be designed whenever possible so the higher pressure and coldest process
stream is on the channel side. This is usually the feed coming into the plant for feed effluent
exchangers. This helps to protect critical parts from sudden changes in temperature. The
hottest part of this stream should be contained in the channel box, leaving the coolest part free
to contact the diaphragm through the hole in the bottom of the partition.
II. Graphite covered CGG (corrugate metal) or KAG (serrated metal) gaskets can be used for tube
sheet and diaphragm gaskets. Clad and double jacketed gaskets are specifically prohibited from
use in screw plug exchangers.
III. The partition assembly shall not be bolted to the tube sheet.
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IV. Tube sheet gasket surfaces must be flat in both the radial and circumferential directions and
perpendicular to the tube sheet.
V. Fo the e uip e t’s
ith MOC of ha
el i te als diffe e t tha that of ha
el heade . It is
recommended to Upgrade of channel internal material to higher alloys to avoid uneven thermal
expansion during the start-up / upset conditions.
6. Conclusion:
Breech lock / Screw plug closures have relatively less weight due to thinner components and less cost
compared to conventional D-type high pressure heat exchangers & conventional flanged channel type of
exchangers, which makes them very popular in refinery applications such as hydrotreaters and
hydrocrackers. Breech lock / screw plug closures have easy assembling/disassembling, lesser number of
flanged joints and very high reliable sealing features thereby minimizing risk of explosion hazard. Based
on past experience of jamming and failure of some internal components and gaskets, manufacturers and
licensors of these closures have modified the design and metallurgy of internals and gaskets. It makes
them more reliable, safe and easy to maintain.
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Prediction of Reformate Octane Barrels for change in Crude Mix using Simulation
Model at HPCL-Mumbai Refinery (MR)
Rohitashva Tewari ([email protected] ), Ashok Golekar ([email protected]),
Ashok Kumar ([email protected] )
1.0
ABSTRACT
HPCL- Mumbai refinery has UOP CCR Platforming unit with an installed capacity of
0.545 MMTPA. Conventionally, the CCR Feed Napthenes and Aromatics (N+2A)
was being monitored and the recommendation for change in Weighted Average Inlet
temperature (WAIT) was given based on the same. This had a delayed effect on
process conditions, resulting in fluctuation in Reformate RON. This had further
impact on Naphtha back-blending, MS production. In order to achieve and sustain the
target Reformate Octane Barrels, a simulation model of CCR unit was developed and
was calibrated with the actual operating conditions in CCR unit. The simulation
model was developed to predict estimated WAIT, gas to oil ratio for achieving the
Reformate Octane barrels along with estimation of Spent Catalyst Carbon. Based on
the simulation model, suggested operating conditions were given to Operations well
in advance for various type of crudes to be processed in HPCL Mumbai Refinery. The
simulation was also used for analyzing sensitivity of Reformate RON, CCR yields
and coke builtup with Gas to Oil ratio.
This facilitated in proper WAIT optimization in CCR unit with the change in Crude
Mix in CDUs. This resulted in 4.7 % higher Reformate Octane Barrels (2016-17 vs
2015-16) on consistent basis after the adoption of the simulation model and preemptive approach of optimization of operating conditions with change in crude mix.
This paper describes the methodology adopted in developing the simulation model to
predict the operating conditions in CCR unit for achieving the target Reformate
Octane Barrels for various Crude mix on sustained basis.
2.0
INTRODUCTION
a.
Reformate Octane Barrels:
The main product of CCR unit is Reformate and the key parameter of this Reformate
is Research Octane Number (RON). In order to meet the Euro IV MS blend RON of
91, the major contributor is Reformate blend stream from CCR unit. The definition of
Reformate Octane Barrels is:
Reformate Octane Barrels = Reformate quantity in barrels * Research Octane Number
of the Reformate
The contribution to MS Blend pool is Reformate Octane Barrels and in order to
achieve the target / maximize the Reformate Octane Barrels either or both Reformate
quantity or Reformate Octane is to be maximized.
b.
Major Variables in CCR Reactor Section:
1
S.No.
1
2
3
4
5
6
7
Independent Variables
Catalyst Type
Reactor Temperature
Space Velocity
Reactor Pressure
H2/HC Ratio
Charge Stock Properties
Feed Additives
Dependent Variables
Catalyst Activity
Reactor Effluent Yields
Product Quality
Catalyst Stability
i.
Reactor Temperature:
Reactor temperature can be defined in two fashions , either Weighted Average
Inlet Temperature (WAIT) or Weighted Average Bed Temperature (WABT). These
can be calculated as follows:
WAIT =
The summation of the
Wt. Fraction catalyst in bed multiplied by
WABT =
The summation of the
Wt. Fraction catalyst in bed multiplied by
c.
Bed Inlet Temperature
Average of the Bed Inlet &
Outlet Temperatures
ii.
Space Velocity:
Space Velocity is a measure of the amount of naphtha which is processed over
a given amount of catalyst over a set length of time. When the hourly volume
charge rate of naphtha and the volume of catalyst are used the term is Liquid
Hourly Space Velocity. (LHSV) The higher the space velocity (lower
residence time) the lower the product RONC or the less the amount of reaction
that occurs at a fixed WAIT. Increased Reactor temperatures will offset this
effect.
iii.
Hydrogen / Hydrocarbon Ratio:
Hydrogen / Hydrocarbon ratio is defined as the moles of recycle hydrogen per
mole of naphtha charged to the unit. Recycle Hydrogen is necessary in the
Platforming Reaction for purposes of catalyst coking rate. It has the effect of
sweeping the reaction product and condensable materials from the catalyst and
supplying the catalyst with readily available hydrogen. An increase in H2/HC
ratio will move the naphtha through the reactor at a faster rate and supply a
greater heat sink for the endothermic heat of reaction. The end result is
decreased catalyst coking rate with little effect on the product quality or
yields.
Estimation of WAIT for achieving target Reformate RON:
The following steps are to be followed for estimation of WAIT in order to achieve the
target Reformate RON with the actual Feed Quality (P (Parrafins) O (Olefins)
N(Napthenes) A (Aromatics)).
i.
ii.
List – RONC, Actual N and A, N+3.5A, LHSV and Catalyst Type
Estimate SOR WAIT at 1 LHSV – Basis Figure I below
2
iii.
iv.
Estimate Correction for actual LHSV – Basis Figure II below
Add the results of Points ii and iii to give SOR WAIT
3
4
3.0
Background:
The unit is running since May 2009. The Reformate is the main contributor blend
stream in Euro III / IV MS Blend for its RON. The CCR Feed sample analysis
(PONA) is done daily on morning CCR Feed sample. So basis the conventional
methodology of WAIT estimation for target Reformate RON (stated above), the
WAIT estimation was done manually everyday and the target WAIT along with other
operating conditions were asked to be maintained. However, this methodology was on
post facto basis (target WAIT based on one day older feed sample). So in order to
achieve maximization of Reformate Octane Barrels, a pre-emptive approach was
adopted viz.
1. Estimation of N+2A of the potential CCR Feed for all the potential Crude Mix
to be processed in HPCL MR as per plan.
2. Based on N+2A and planned CCR Feed rate, estimation of target WAIT for
achieving target Reformate RON through a rigorous excel based model.
3. The calculation of estimated WAIT and other operating conditions through a
CCR Simulation Model with the objective of maximization of Reformate
RON and Reformate Yield
4. Information to Operations & Planning for the target WAIT required to achieve
target Reformate RON for all the Crude Mix in advance. Also, considering the
unit constraints, the feasible WAIT for all the Crude Mix.
The above approach helped us in achieving 11.6% more Reformate Octane Barrels
compared to the previous conventional approach.
The paper below highlights the following steps taken to adopt the approach stated
above for maximization of Reformate Octane Barrels:
a.
b.
c.
d.
e.
4.0
Development of excel based rigorous model and estimation of target WAIT
corresponding to N+2A for CCR Feed of various Crude Mix.
Development of Simulation Model and WAIT estimation along with
sensitivity analysis done in Simulation Model
Communication to Operations and Operations Planning for the estimated
WAIT along with other operating conditions for all the potential Crude Mix
Improvement achieved in Reformate Octane Barrels
Summary
Estimation of WAIT in CCR for various Crude Mix:
The PONA (Parrafins, Olefins, Napthenes & Aromatics) of the CCR Feed were
estimated for neat 100% processing of Mumbai High, Arab Extra Light, Murban,
Basrah Light, Upper Zakhum and Das Blend basis the Crude Assay Test run reports
by HPCL R&D.
An excel based model based on R264 catalyst type, actual catalyst distribution in
Reactor (wt%), feed rate of 95 m3/hr, 100 m3/hr and 105 m3/hr and the correlations
of LHSV for WAIT estimation was developed to calculate the estimated WAIT for
achieving the target Reformate RON.
5
After the inputs of CCR Feed PONA, target Reformate RON, correction factor for
actual LHSV and catalyst type were inserted into the manual the estimated WAIT was
calculated using the principles in Operating Manual however with more accuracy and
faster speed.
5.0
CCR Simulation Model development with WAIT estimation and Sensitivity
Analysis:
The CCR Simulation Model was earlier built by HPCL R&D in 2012, however, the
same was not being used and was required to be recalibrated. With the help of M/s
ApenTech, the CCR Simulation Model was recalibrated to actual operating conditions
and R264 catalyst. Subsequently the model was used for WAIT estimation to achieve
target Reformate RON.
i.
Estimation of WAIT for target Reformate RON:
6
The RON increases with increase in WAIT. The carbon compounds that have high
RON (Aromatics and Branched Naphthenes) are products of reactions which are
endothermic in nature (Dehydrogenation, Cyclization). Since endothermic
reactions are favoured by high temperatures, hence production of high RON
carbon compounds increases as WAIT increases, thus increasing the RON.
So with the given CCR Feed rate and N+2A, the estimated WAIT for achieving
the Reformate RON was estimated. Also, the sensitivity of Reformate RON with
the WAIT was also studied.
ii.
Sensitivity of Reformate RON with Gas to Oil Ratio:
7
The RON decreases as H2/HC Ratio increases. Reactions producing high RON
carbon compounds have H2 as a product. High H2/HC Ratio drives these reaction in
the backward direction according to Le Chatelier‘s Principle. This leads to decreased
production of high RON carbon compounds, hence decreasing the RON.
iii.
Sensitivity of Reformate Yield with WAIT:
8
1) The Reformate yield (both in LV% and Wt%) decreases with increase in WAIT. This
is because as temperature increases, the conversion due to hydrocracking reaction
increases as it is a high activation energy reaction. This leads to cracking of higher
carbon number compounds into lower carbon number compounds, thus reducing the
Reformate Yield (C5+ components).
2) The Reformate yield decreases as H2/HC Ratio increases but the change is not so
significant. The increase in the amount of H2 increases the conversion of the
hydrocracking reaction as hydrogen is a reactant in the cracking reaction and this
drives the reaction in the forward direction according to Le Chatelier‘s Principle.
9
iv.
6.
Sensitivity of Coke Laydown on Catalyst with Gas to Oil Ratio:
1.
The coke laydown on catalyst increases with increase in WAIT. This is because
higher temperatures favour the coking reactions.
2.
The coke laydown on catalyst decreases as H2/HC Ratio increases. H2 helps in
sweeping of the precursors that lead to coke formation. Hence higher amounts of
H2 reduces coke laydown on catalyst.
Communication to Operations & Planning for Operating Conditions:
As a pre-emptive approach, the estimated WAIT for achieving target Reformate RON
for all the Crude Mix along with the operating conditions was communicated to
Operations and Operations Planning.
N+2A, vol%
WAIT reqd for
RON of 102,
deg C
Feasible WAIT,
deg C
Feasible
Reformate
RON @ feed
rate of 95 m3/hr
Feasible
Reformate
100%
100%
Mumba Arab
i high
Extra
Light
73.9
45.5
100%
Murban
100%
Basrah
Light
100%
Upper
Zakhum
100%
Das
Blend
44.3
56.8
30.0
53.5
532
548
548
542
556
543
538
538
538
538
538
538
105.1
98.5
98.4
101
95
100.3
104.6
98
98
100.4
94.4
99.8
10
100%
100%
Mumba Arab
i high
Extra
Light
RON @
rate of
m3/hr
Feasible
Reformate
RON @
rate of
m3/hr
100%
Murban
100%
Basrah
Light
100%
Upper
Zakhum
97.4
99.9
93.9
100%
Das
Blend
feed
100
feed
105
104.1
97.5
99.3
The above inputs helped in Operations for maintaining the required operating
conditions with change in Crude Mix and Operating Planning for optimum planning
of crude processing from Reformate RON point of view.
7.
Improvement achieved in Reformate Octane Barrels:
Reformate
(TMT)
RON (201617)
RON (201516)
8.
508.5
%
increase
in
Octane
Reformate Million Reformate
Barrels
Octane Barrels
391.5
4.7
479.90
373.9
Summary:
The pre-emptive approach of using simulation model for estimating operating
conditions in order to achieve target Reformate Octane Barrels and operating
conditions helped us in achieving 4.7% higher Reformate Octane Barrels in 2016-17
vs 2015-16 with reduction in Specific Fuel Consumption by 9.7% worth equivalent
4000 SRFT.
11
OPTIMISING FLOODED COLUMNS -PRACTICAL EXPERIENCE
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P.
KAMATHPROCESS ENGINEERING DEPT., MRPL
ABSTRACT
Key
Word:
In most Fluidized Catalytic Cracking Units
Debutanizer, Flooding, ABCD, Gamma Scanning
PFCC,
Absorber,
Stripper,
flooding is frequently observed in one or more of
their columns restricting the product recovery
quality and processing rate, there by affecting the
profitability of the plant to a great extent. The
flooding can be caused by poor liquid distribution,
mechanical damage to trays, or change in feed
composition.
Diagnosing and eliminating the root cause is
of prime importance to resolve the issue. In case of
hardware limitation the column can still perform
satisfactorily with modified operating conditions
till the expected turnaround of the plant. But in case
the flooding is caused by feed composition change,
column parameters are to be adjusted to overcome
the flooding condition and to bring the column
back to a stable condition. Flooding in a column
can be identified by rise in differential pressure
across the column, high liquid entrainment from
top of the column, reduction in column outflows or
Introduction
PFCC
unit
employ
deep
catalytic
cracking
technology, the catalytic cracking is achieved
through contacting feed with fluidized catalyst in
reactor riser. Products are separated to the desired
purity in the downstream Gascon section. Gascon
section consists of Main Fractionator (MF), Wet
Gas
Compressor,
High
Pressure
Separator,
Absorber, Stripper, Debutanizer and Depropanizer
sections. Gas from MF overhead is compressed to
17 KSC by WGC and the hydrocarbon is separated
in HP Separator after cooling down to 40 OC by
HPS Feed Condenser. The gas fraction from HPS is
then routed to Absorber to recover Propylene by
absorbing with supplementary lean oil. Liquid
hydrocarbon from HPS is routed to Stripper. In
Stripper lighters from the liquid stream are stripped
out using LCO & LP Steam reboilers.
sharp change in column temperature profile.
As a whole Absorber, Stripper and HP Separator
This paper intends to discuss experience of
work in tandem to produce Dry gas from the
flooding problems occurred in Stripper, Absorber
Absorber top, and liquid hydrocarbon consisting of
and Depropanizer, behaviors of columns during
LPG, Propylene, and Naphtha Stream from Stripper
flooding and related recovery issues in Deep
bottom. The objective of this system is to minimize
Fluidized Catalytic Cracking Unit. An Algorithmic
Propylene in Dry gas and Lighters in Debutanizer
Basic Cause Detection (ABCD) chart is established
feed stream. The efficiency of this system
to find the possible root causes of the flooding
determines the efficiency of the unit.
problem caused by compositional changes, using
the available process parameters. Gamma scanning
Flooding
of column was carried out to identify inception of
column of PFCC is observed at different operating
flooding in the Stripper column. With the help of
conditions and time leading to recovery issues.
UNISIM process simulation software, the column
was simulated to find the optimum parameters
required to normalize the column. The conditions
used in the algorithm are successfully tested
manually during the flooding condition.
in
Stripper,
Absorber,
Debutanizer
Use of algorithmic method: Algorithmic Basic
Cause Detection (ABCD) and UNISIM process
simulator is shown to stabilization Different
column.
OPTIMISING FLOODED COLUMNS -PRACTICAL EXPERIENCE
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P.
KAMATHPROCESS ENGINEERING DEPT., MRPL
BACKGROUND
190MT with increasing HPS temperature from 37
Behavior of Flooded column
to 67 Deg C.
Absorber Flooding: It is observed that HPS feed
0,13
0,50
0,12
condenser exchangers developed tube leaks calling
for isolation of the exchangers on cooling water
side. The isolation of exchangers on cooling water
0,45
0,11
0,10
0,40
0,09
side has increased the stripper feed temperature
from 40OC to 65OC Since load on Absorber went
up flooding started in Absorber because of which
C3+ in Abs Feed, mol
fraction primary axis
C2- in Stripper Feed
0,35
0,07
0,06
0,30
Plant load could not be increased beyond 80%.
0,08
550
37 39 42 46 50 54 58 64 68
230
500
180
It was also observed that behaviour of absorber
column was very sensitive during that period and
even slight change in flow to Absorber caused by
change in carrier gas flow, Dryer or RSH-COS bed
depressurization to MF distillate drum or even
Stripper Feed,TPH
change in ambient temperature was causing the
Absorber feed TPH
Absorber column flood.
number
of
80
37 39 42 46 50 54 58 64 68
Fig1: Lighter shifting from Stripper to Absorber
400
While Absorber flooding, it is observed that
maximum
130
450
times
initially
Upper
Feed v/s HPS Tem.
Absorber Reflux flow increased drastically from
290m3/hr to 380m3/hr caused high DP in
After rectifying HPS Condenser tube leak in
Absorber. Later the flooding travelling up in the
Shutdown and taking in line HPS condenser,
column to the reflux drums causing high DP across
absorber problem resolved completely.
the column.
During Absorber flooding it is observed that
Propylene slipping into dry gas increased more
than the design allowed figures even though
supplementary lean oil is on the higher side
suggesting poor performance of the trays during
flooding time. The phenomenon was simulated in
UNISIM, which shows both feed rate and C3 load
in absorber increased drastically which start
causing Absorber flooding. As shown in curve
Stripper load reduced from 530MT to 440 MT
where as Absorber load increased from 100 to
Depropanizer Flooding
For Certain period continuous Flooding conditions
were observed in the Depropanizer column
rectification section as Delta pressure reached up to
0.8 Kg/cm2. Decrease in efficiency of the
Depropanizer Column tray was observed as In
flooding condition C4 in PRU feed which is the top
product in the Depropanizer column used to go as
high as 4~5wt% against the design of 0.5wt%,
propylene slip from Depropanizer bottom is as high
as 9~10wt%. Depropanizer top temperature used to
go as high as 50OC against the design of 46OC due
to heavier component accumulation. Increase in
OPTIMISING FLOODED COLUMNS -PRACTICAL EXPERIENCE
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P.
KAMATHPROCESS ENGINEERING DEPT., MRPL
reflux flow had very little impact on top
Debutanizer column or Caustic Strength failure;
temperature control.
this also limits the overall plant throughput.
350
0,6
300
0,5
250
0,4
Case of Flooding:
Algorithmic Basic Cause Detection (ABCD) chart
was established to find the possible root causes of
200
0,3
150
100
Avg Reflux
50
Avg DP
0
the flooding problem caused by compositional
0,2
changes, using the available process parameters. It
0,1
is a comparative method which utilizes the Normal
0
column (with similar load) data along with the
Fig 2: High Delta Pressure with Reflux
Flooding
data
using
simple
technique
like
Important Parameter Variation, temperature profile
Based on Primary analysis caustic carryover was
across the column, L/V ratio, etc. It determines the
suspected and same is supported by high Delta
most possible cause of flooding and possible
pressure across the treated LPG Coalescer and
column parameters that need to be adjusted to
filter.
overcome the flooding condition and to bring the
Later Hot water wash of column was
conducted which resolved the Column high Delta
column back to a stable condition.
pressure issues also Red oil formation control steps
were taken, which had affected the performance of
LPG Coalescer.
Vapour entrapment caused DP: It is observed
that with change in MF OVHD Flow at certain
stripper column feed rate, column Delta Pressure
Stripper Flooding: It is observed that Stripper
changes. With increase in MF overhead flow
column operation is very sensitive and affect by
Stripper OVHD Vapour flow decreases and
various factors like feed composition, Reboiling
Stripper bottom Draw increases causing vapour
efficiency and Supplementary lean oil rate, Feed
entrapment and Delta pressure across the column.
water content etc.
As
MF
overhead
flow
is
heavier
than
Supplementary lean oil also contains minor water,
High Delta pressure across the column is primary
sign of Flooding, during the flooding fuel gas may
have Propylene 10-12% also Slippage of lighters
(C2/H2S) to Downstream Unit is possible. High
slippage for long time can lead to Pressurization of
increase in liquid flow and decrease in L/V
(wt%/wt %) ratio is observed. In Such cases
Column
is
stabilized
by
reducing
the
Supplementary lean oil and overall feed to stripper
column also reboiling was adjusted accordingly.
OPTIMISING FLOODED COLUMNS -PRACTICAL EXPERIENCE
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P.
KAMATHPROCESS ENGINEERING DEPT., MRPL
1,2
40,0
1,0
30,0
0,8
0,4
0,0
causing Delta pressure across the column Gamma
Scanning was conducted to check the column
20,0
0,6
0,2
Since with the small change in feed or reboiling is
STRIPPER DP
MF OVHD DIST OUT m3/h
1,2
1,0
0,8
0,6
0,4
0,2
STRIPPER DP
STRIPPER FEED
0,0
10,0
0,0
700,0
680,0
660,0
640,0
620,0
600,0
580,0
560,0
internals.
Gamma Scanning of Stripper column: Gamma
Scanning is a Sealed Source Radiation Technique.
Sealed source radiation (Cobalt 60, λ-5.27 yr,
Intensity 80mCi) were used to penetrate the
Stripper column. When  - ray passes through
various medium from the source to detector, some
of its radiation is absorbed by medium. The amount
of radiation that is not absorbed is given by [1]
I = I0 e -x
Fig4: Vapour entrapment caused DP
Reboiling induced Delta Pressure: At constant feed
rate with increment in Reboiling, DP is observed
which lead to similar vapour entrapment and increase
Where:
I = Radiation intensity in counts per second, as seen
by
the
detector
Io = Intensity of radiation from the  - ray source
reaching
the
detector
in
the absence of the material in counts per second
 = Linear absorption coefficient (cm –1)
in bottom draw, with reduction in overall feed stripper
x=Thickness of material
DP is normalized
=Density of medium
1,0
104,0
0,8
103,0
A plot is drawn between the height and counts
0,6
provide internal view of Stripper column.
102,0
0,4
0,2
0,0
DP
Bottom tem
101,0
100,0
1,0
500,0
0,8
490,0
0,6
480,0
0,4
470,0
0,2
DP
460,0
0,0
Stripper Feed
450,0
Fig5: Reboiling induced Delta Pressure
Fig 6: Gamma Scanning Pattern
OPTIMISING FLOODED COLUMNS -PRACTICAL EXPERIENCE
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P.
KAMATHPROCESS ENGINEERING DEPT., MRPL
Inferences of Gamma Scanning: All trays of
stripper column were found intact. Flooding was
50
40
observed mostly in the below chimney tray (5-39).
This result was farther analyzed with the help of
Simulation.
30
Normal Column
Flooded Column
20
Simulation Result of Stripper Column: Stripper
10
top pressure is determined by Delta Pressure across
0
the column, for Different top pressure, bottom
45
55
65
75
85
95
temperature that is needed is different. Also for
different feed composition required Stripper bottom
Actions in Pipe line to rectify Stripper DP
temp is different.

116
Advance Process control for Stripper
Column: Since various factors affect the
performance of the column. It is very
114
important to take the action in very
112
advance and APC implementation will
Stripper Bottom
Temperature OC
110
help to achieve the same.

108
15,00 15,50 16,00 16,50 17,00 17,50 18,00
C2 and H2S analyzer to achieve the
optimum reboiling by tuning the C2 and
H2S escape from bottom.
Fig7: Variation of Bottom Temp v/s Top Pr.

leveling and down comer configuration in
Variation in temperature Profile in flooded or
next possible opportunity.
normal Column operation is not very much
Pronounced, This may be due to close boiling of
Column detailed inspection including tray
Conclusion
the mixture at Stripper condition. However the
Efficient Column operation is extremely important
Drop of temperature across the 5 to 35 tray shows
to harness the best from any plant. Tools like
below chimney tray as flooding, top and Bottom
Algorithmic cause detection, simulation are really
temp
helpful for immediate actions to stabilize the
affects
due
to
Reboiling
and
Feed
Temperature however variation in below chimney
column.
tray is mostly due to abnormality in mass transfer.
Reference
50
1.
40
Henry Z. Kister, Distillation Operation,
30
Ch14,
20
Techniques.
10
0
45
55
65
75
85
Simulation Tem Profile
95
105
Radioisotope
Troubleshooting
105
Approach towards Abatement of VOC Emission
AT Paradip Refinery
J R Behera (DM-HSE), Shashi Vardhan (DGM-TS)
Introduction:
The petroleum refining industry converts crude oil into petroleum derivatives such as liquefied
petroleum gas, gasoline, kerosene, aviation fuel, diesel fuel, fuel oils, lubricating oils, and feed
stocks for the petrochemical industry. Petroleum refinery activities start with receipt of crude
for storage, include all petroleum handling and refining operations and terminate with storage
prior to shipping the refined products from the refinery. During these operations, waste water
collection and treatment system becomes the part of the operations of Petroleum Refineries.
Because of all these activities, Refinery operation encounters with fugitive emissions.
Fugitive emissions in petroleum refining facilities may escape from leaking tubing, valves,
connections, flanges, gaskets, steam traps, packing, open-ended lines, floating roof storage
tanks and pump seals, gas conveyance systems, compressor seals, pressure relief valves,
breathing valves, tanks or open pits/containments, oil-water separators, and storage, loading
and unloading operations of hydrocarbons. Depending on the refinery process scheme, fugitive
emissions may comprise:





Hydrogen
Methane
Volatile organic compounds (VOCs), (e.g. ethane, ethylene, propane, propylene,
butanes, butylenes, pentanes, pentenes, C6-C9 alkylate, benzene, toluene, xylenes,
phenol, and C9 aromatics)
Polycyclic aromatic hydrocarbons (PAHs) and other semi-VOCs
Inorganic gases, H2S, ammonia (NH3), CO, CO2, SO2 and sulfur trioxide (SO3) from
sulfuric acid regeneration in the sulfuric acid alkylation process, NOX etc
This paper will briefly describe the VOC emission and its control measure in the following two
areas of operation in Paradip Refinery:
1. Product Loading in Refinery owned South Oil Jetty (SOJ)
2. Effluent Treatment Plant of Paradip Refinery (ETP)
What is VOC:
Volatile Organic Compounds (VOC) are organic chemicals that easily vaporize at normal
conditions and enter into the atmosphere. They are organic as they contain carbon atoms in
their molecular structure. VOC may include a very wide range of individual substances, such
as hydrocarbons (eg methane, ethane, benzene, toluene, etc.), oxidized hydrocarbons (or fuel
oxygenates, such as methyl tert-butyl ether (MTBE)) and by-product organic compounds from
chlorination in water treatment (such as chloroform).
Harmful Effect of VOC emission:

The VOC given off through vaporization of crude oil and refined products are a mixture
of light-end hydrocarbon components such as methane, ethane, propane and several
other gases. Methane, being lighter than air, will be emitted to the atmosphere and
contribute to the greenhouse effect.
 All the other components generally referred to as non-methane VOC (NMVOC), being
heavier than air, will react on warm days with nitrogen oxides (NOX) in the air and
form the ground level ozone commonly known as smog. The ground level ozone has a
detrimental effect on human health, vegetation and buildings.
 Some NMVOC are quite harmful, including benzene, toluene and xylene, which may
cause leukemia.
So, to keep environment safe, it is required to adopt technology for effectively control of VOC
emission.
Approach of Paradip Refinery towards abatement of VOC emission:
1.
Vapor Recovery Unit at South Oil Jetty (SOJ)
Paradip Refinery Project, IndianOil, has set up the
Oil Jetty along with associated facilities at Paradip
Port Trust. A single berth has been provided for
loading of products and unloading of crude oil.
Following petroleum products are dispatched
through Jetty:
SL NO
Product
Crude
1
Propylene
2
Propylene-vapor return
3
MS-Regular/ Premium
4
Naphtha
5
ATF
6
HSD
7
Among the above products loading, vapour
recovery system has been installed for Naphtha and
MS.
Product Loading at South Oil Jetty (SOJ)
Vapor Recovery Unit at South Oil Jetty (SOJ)
VOC Recovery:
Objective:
Vapor Recovery System at SOJ installed by Paradip Refinery is to adhere to the environmental
emission norms during ship loading because volatile organic Compounds (VOC) emissions
resulting from the loading of Naphtha and MS have a very negative impact on the environment
like:
 Human health
 Pollution of the troposphere
 Jetty Safety
Working Principle:
1 - Adsorption of the hydrocarbons on activated carbon
2 - Regeneration of the carbon by means of vacuum by dry vacuum pumps
3 - Re-absorption of the hydrocarbons in a liquid product
Process Description:
Vapour collection system
The mixture of hydrocarbons and air present inside the storage to be loaded is displaced by the
liquid. The net pressure required to transfer the mixture to the VRU is limited by the maximum
allowable pressure in the storage (limited by the settings of the pressure safety devices).
Adsorption cycle
Vapor Inlet
Clean Air outlet
The adsorption process is discontinuous (adsorption
cycle then desorption/regeneration cycle). To
guarantee a continuous process, the VRU is
equipped of pairs of adsorber vessels. After a certain
period of time, one bed is saturated with
hydrocarbons and the vapours flow is switched to
the other bed, whilst the first bed will be regenerated
by pulling vacuum. In the adsorption process, the
hydrocarbon/air mixture gets through the bed of activated carbon where the hydrocarbons are
adsorbed on the internal surface of the activated carbon. The purified air leaves the adsorber
vessel through valve and the outlet vent stack.
Desorption cycle
To regenerate the adsorber vessel in-line, vacuum pumps pull vacuum in the vessel ready for
regeneration. By lowering the pressure, the
process of adsorption is reversed and the
hydrocarbons leave the surface of the
activated carbon and transferred by the
Vessel ready
Vessel ready for
vacuum pump to the re-absorber column. If
for adsorption
regeneration
the pressure into the adsorber vessel in
regeneration is sufficiently low, the purge
valve is opened to clean the bottom layer of
activated carbon with outside air. After the
purge phase (around 5 min at the end of
desorption cycle)), the valves are closed and
the regeneration of adsorber vessel is
finished.
Absorber Column:
The vapours, leaving the vacuum system, flow upwards through an absorber column. They are
brought in contact with a counter flow of absorbent. A small amount of air saturated with
hydrocarbon vapours leaves the absorber column from the top and fed back in the inlet of the
VRU. The absorbent for the MS and Naphtha
is same liquid hydrocarbon. This comes from
the product line itself. The absorbent is injected
at the top of the absorber column and
distributed equally over a bed of packing rings.
In the sump of the absorber column, the
absorbent and the recovered hydrocarbons are
accumulated and pumped back to the same
product line by absorbent return pumps.
2.
Vapor Capture system in Effluent Treatment Plant ( ETP) of Paradip refinery:
In Paradip Refinery, to treat waste water, a state-of-the-art ETP with VOC capture system has been
installed. The typical configuration of ETP of
Paradip Refinery consists of:
Waste water collection systemAPI/TPI
Equalization tank Oily DAF Bio Tower 
Aeration Tank  Clarifier  Bio DAF  Final
collection Basin. Oil recovered in API/TPI is
collected in Slop tank. VOC system is installed all
the oil laden sources like Slop Tank, Brine Tank,
Equalization Tank and Oily DAF.
3.
Vapor Capture:
Objective:
Waste water generated in the petroleum
refinery during the refining of crude oil and
other auxiliary activities has a significant
quantities of hydrocarbon which in turn
becomes a major source of VOC emission
which impact on environment. To contain this
VOC emission, Paradip refinery has installed
VOC capture system in its state-of-the-art
ETP.
Volatile Organic compound (VOC) generation is from the following sources:
 Slop tank,
 Equalization Tank,
 API/TPI/Oily DAF,
 Spent caustic units tanks
Working Principle:
VOC laden gas from the above sources is
collected in a common header and passed
through the carbon bed. The carbon bed
adsorbs these VOCs as they pass through
the bed and clean gas sent to stack.
Process Description:
Each source is connected to a carbon bed system. Two sets of carbon beds are used to serve all
sources, each with a primary carbon canister and secondary carbon canisters used to back up
the primary canisters.
The fume from the sources contains
organic vapors including volatile and
semi-volatile organic compounds,
referred as VOCs. The carbon bed
adsorbs these VOCs as they pass
through the bed. The gas that exits the
bed is moist air, with small amounts of
VOC, CO, and non adsorbable organics
such as methane. With long term use, the bed becomes saturated with organics and can no
longer adsorb VOCs. This is when VOC "breakthrough" occurs.
The carbon beds are equipped with up and downstream thermocouples and a water deluge. The
deluge is activated when a temperature rise across the beds is excessive. If unchecked, the bed
would ignite. The high temp limit first turns off the ID fans, shuts the system down, and closes
a damper to the beds and introduces N2 purge to canisters. If the temperature continues to rise
the water deluge is activated. A water valve is opened introducing plant water at pressure into
the canisters
Conclusion:
It has been discussed in the paper on the harmful effect of VOC emission from Refinery
operation starting from crude import to product loading. So keeping in view of the environment
impact as well as to adhere to environmental emission norm, Paradip refinery has taken
proactive step from the inception itself. Therefore, in the design itself, VRU has been installed
in South oil Jetty (SOJ) for Vapor Recovery and Vapor Capture system in the waste water
treatment section i.e in Effluent Treatment Plant (ETP). By this, Refinery will ensure reduction
in VOC emission during its operation.
***********
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
-
Pranjal Changmai, DGM (Chemical)
Pradeep Rawat, Chief Manager (Chemical)
Pranjal Kumar Phukan Sr. Manager (C&P)
Subodh Kumar, Dy. Manager (Chemical)
Tamagna Ghosh, Dy. Manager (Chemical)
Manish Kumar Binjola, SO (C&P)
Abstract:
Setting up a petrochemical plant and maintaining the operations is a challenging task considering
different factors like demographic position, climatic conditions and logistic issues. In view of spiraling
availability of the crude oil, associated gas and Refineries, it has become almost inevitable to look for
value added opportunities to be integrated with the already existing upstream facilities through
Petrochemical integration. The fact that most of the petrochemical produces, invite a higher degree of
margin vis-à-vis the fuels, there is a strong case for integration between these Complexes involved in
Exploration/Refinery and Petrochemical Complex, wherein, both feed as well as energy integration can
be exploited for soliciting higher revenues.
In the course of the study of various petrochemical industries, it is being found that generally
petrochemical industries are located nearer to sea-ports for facilitating coordinated and planned
movement for both inward and outward transportation.NE India is already lagging to have an
inherent disadvantage of being away from ports hence, the supply of plastic materials from other
parts of the country become costly affair to this remote location. However, this same situation
makes the domestic plastic processing more competitive and provides significant opportunity at
this place.
The per capita consumption of plastic in the North-Eastern region is ~2.8 kilogram against a
national average per capita consumption of ~9.7 kilogram. The relative global consumption
figure is around ~27 kilogram. Until recently, the entire North-Eastern region including Eastern
region accounted for a share of approx. 9 per cent of the polymer consumption in the country
though North-Eastern region has the share of 3.8 per cent of the population. With the set-up
of Brahmaputra Cracker & Polymer Limited (BCPL), this region will get the necessary fillip in
feedstock sourcing and in the polymer processing sector along with increased growth in Enduser industries which could propel the growth of plastics in North-East India further.
For any petrochemical industries, supply chain is complicated activities due to limited
availability and hazardous nature of chemicals. In the context of the industry referring in this
case study it is noted that the plant is located in a remote location far-off from any of the Indian
sea-ports and having poor road infrastructure leading to imprecise shipping accuracy, frequent
network designing in accurate supply chain time frame. Most of the chemicals consumed are
hazardous in nature which requires proper and controlled on-time fulfillment, storages, and
preservation till distribution to the plant.
Key words: Petrochemical, Chemicals, Hazardous, Per capita consumption, Remote
location, Inland transport, Ports, Logistics
P a g e 1 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Brief History:
The Assam Gas Cracker Project was proposed as a part of the implementation of Assam Accord
signed on 15th August, 1985.A letter of intent was issued in favor of Assam Industrial
development Corporation (AIDC on 25.01.1991 to set up a Petrochemical Gas Cracker Complex
of 3 lakh tons per annum Ethylene capacity in the joint sector. Due to inherent disadvantages of
locating a project in Assam, Government of Assam (GoA) could not attract a joint venture
partner in spite of their best efforts. Therefore, Assam Government approached Central
Government in April, 1992 for fiscal concessions in order to neutralize the disadvantages of
setting up the project in Assam.
Accordingly, Memorandum of Understanding (MOU) was signed between the Government of
Assam and Reliance Industries Ltd. (RIL) on 20th May 1994 for setting up of this project and
Reliance Assam Petrochemicals Ltd. (RAPL) was formed for implementing the project. The
shareholding Pattern Envisaged was 11% Assam Industrial Development Corporation (AIDC),
40% RIL and 49% by the public. However, RAPL did not start work for implementation of the
project due to non-availability of sufficient natural gas to produce 2 LTPA ethylene. RAPL
demanded that the shortfall in natural gas to be compensated by supply of LPG at natural gas
price and the project remained a non-starter.
But later on, in a meeting held on 20.02.2003 under the Chairmanship of Additional Secretary,
Department of Expenditure, Ministry of finance, to discuss the issue of subsidy for the associated
gas for the Assam Gas Cracker project, it was decided that GAIL would examine the feasibility
of taking up the project on its own. GAIL would also indicate the assistance needed from other
PSU’s and the support needed from Government of India for setting up the project.
Finally, Cabinet Committee on Economic Affairs (CCEA) in its meeting held on 18th April 2006
accorded approval for setting up of Assam Gas Cracker Project at Lepetkata, District Dibrugarh,
Assam. The Joint Venture Company in the name of BCPL was incorporated on 8th January
2007.
KEY INFORMATIONS
Project commencement date: 09.04.2007
Total land area at Lepetkata: 1181 acres
Initial Approved Project Cost: Rs. 5461 Cr.
Final Approved Project Cost: Rs. 9965 Cr.
Dedi ation to the Nation y Hon’ le
Prime Minister – 05.02.2016
P a g e 2 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Introduction:
Brahmaputra Cracker and Polymer Limited, a Petrochemical Complex is located at Lepetkata.
The principal end products of the complex are High Density Polyethylene (HDPE) and Linear
Low Density Polyethylene (LLDPE) totaling 2, 20,000 Tonnes per Annum (TPA) and 60,000
TPA of Poly-Propylene (PP). The other products include Hydrogenated Pyrolysis Gasoline and
Pyrolysis Fuel Oil.
P a g e 3 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Main Body:
Taking into account the risks associated with setting up of a green field mega project in Assam, it
is imperative that the viability of the project be sufficiently ensured. This requires (i) reducing
the level of investment (ii) confirmed feedstock availability at reasonable cost and (iii)
competitive finished product pricing considering remote location of the plant. The petrochemical
business being cyclical in nature, the project financials must be sound enough to absorb the
polymer price fluctuations.
BCPL- A Case study
Natural gas supply commitment from OIL & ONGC, combined with moderate pricing have
given a strong economic advantage to BCPL that turn gas and natural gas liquids into
intermediate chemicals and an array of finished polymer products that are among the most
commonplace manufactured objects in modern life.
The contribution of BCPL share in the existing/future manufacturing capacity is shown in table:
We too believe that the impact of this industry build out will be felt statewide, not just in Northeastern states of India having per capita consumption of 2.8 kg, with a “ripple effect” providing
opportunities for decades to come across the other states of India where per capita consumption
is 9.7 kg.
P a g e 4 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
The BCPL is a first step forward towards the build out of a regional petrochemical hub and
manufacturing renaissance in the plastics and chemicals sector in the North-east that will lead to
new business creation and expansions not just in this sector, but in many sectors that provide
support products and services to it. Both new and existing companies will take advantage of the
polyethylene and polypropylene being produced by BCPL and intermediates produced from
other potential regional cracker facilities. This paper goes beyond simply focusing on the BCPL.
P a g e 5 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
We attempt to define the challenges, experience and innovation during different phases of this
Petrochemical project as known today, and the reason for projected timeline delay in its
construction, as well as change in the estimated value of the project, overall expected build out of
the petrochemical sector in this region. Before detailing on the challenges and innovative
solutions, a brief outline with respect to “Concept to Commissioning” is required.
Major Challenges overcome during initial and construction phase of the
project:
 Bidders’ response was poor considering the plant capacity leading to delay in finalization of
licensor for ECU. Subsequent to it, the license for downstream units i.e., PE and PP unit
also got delayed. Key time frame was consumed during changes made in detail feasibility
based on BDEP Package received from Licensor.
P a g e 6 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Mitigation plan to bridge this gap/slippage in Schedule:
- Once the Licensors got selected, signing of license agreementwere done at the
earliest.
- Licensors being expedited to issue final datasheets at the earliest.
- Instead of waiting for the Final Process Packages, Procurement / Engineering
activities were commenced based on preliminary datasheets.
- Completion of Site Office, Construction Substations & Receiving Substation
buildings achieved by adopting unconventional construction techniques.
- Basic and Detail Engineering also being executed from EIL, regional office Kolkata
& regional office, Chennai along with EIL, Delhi office parallely for facilities
development at Lakwa, Duliajan & Lepetkata. However major procurement were
executed by EIL HO.
- Process Licensors were pursued to supply critical Equipment and long lead items
Datasheets, P&ID’s, Equipment Layouts etc. within 12-16 weeks as against
Contractual date for supply of BDEP from 26 to 36 weeks.
- Cracker Unit Licensor was also pursued to provide the Heater’s Detailed Engineering
Package within 36 weeks (i.e. Nov’09) against contractual 54 weeks (i.e. Mar’10).
 Delay in permission from statutory authority regarding minor minerals used in the project.
- Persistent persuasion with different statutory body.
 Frequent Theft / sabotage, misplacement of project materials.
- Enhanced security measures have been taken to overcome these issues at the
maximum possible extent.
 Loss of working days due to frequent bandh / strikes.
- Relentless dialogues with Local administration and appropriate manpower planning
with hutments/accommodation arrangement inside the plant premises.
 Site grading & piling work affected due to incessant rain leading to submerging & inherent
high level of underground water.
- To ensure endurance of work and to facilitate smooth movement of Equipment and
Manpower during monsoon respective Site Grading Contractors were advised to
complete the construction activities for “Approach roads” along with Culverts and
Drains at the working sites.
P a g e 7 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
 Delay in readiness of utilities due to lackluster approach of LSTK (Lump sum turn-key)
contractors. Financial Crunch faced by the project executing agencies results in poor
performance.
- Offloading of such contracts including LSTK and job was completed at the risk and
cost of the main contractor.
 Scarcity of water occurred due to disruption of HT cable.
- No. of bore wells increased to fulfill the demand of raw water.
 Paucity of requisite skilled and unskilled manpower.
- Later Skilled and experienced contract man power was outsourced.
 ODC Consignments–transportation of over design consignments (ODC) to site. Delay in
erection of major equipments such as Purge Column, Ethylene fractionators etc.due to poor
road/rail infrastructure and non-availability of Sea port nearby to the Project site.
- Equipments transported in knocked-down state.
- Equipment fabrication at site such as PE Polymerization Reactor, Hydrogen Bullet,
Spheres, Purge Column etc.
- Routes were designed for Road and Water transport.
Transportation by Water Ways: Weight upto 550 MT and length upto 55 M planned
during monsoon season only (July to September).
Selected Route
Kolkata/Haldia--Raimongal—Chalna--Khulna--Mongla--Kaukhali--Barisal--Hilza-Chandpur-- Narayangunj-- Aricha—Sirajgunj--Chilmari--Dhubri –Jogighopa – Tejpur Dibrugarh.
Transportation by Road Ways: Packages having dimension of 30mx5.2mx4.7 –100 MT
safely transported by road.
Selected Route
NH37-Kodarma--Begusarai--Purnia--Dalkola--Siliguri-- Barobisa-- Guwahati—Jorhat--Site
(Dibrugarh).
P a g e 8 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
 Non-availability of high capacity crane, leads to delay in erection of heavy equipment.
The road and site infrastructure was unconducive for its movement.
`-Nevertheless, BCPL developed the site condition on priority basis and hire high
capacity crane from outside agencies and the entire assembly took place at site for
erecting these heavy equipment.
P a g e 9 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
 Likely project time & cost over-run
-
The projected budget and timeframe for building the entire plant including all
supporting units was projected Rs. 5461 crores, for a span of 5 years but later on
vary substantially due to high escalation rate (global recession) during 2008-2010,
prolonged monsoon and many other factors arises during the initial and final
stages of the project results in final project cost to Rs. 9965 crores.
 Dearth of industrial Culture & awareness/aptitude among people in the region about the
benefits of Petrochemical products.
- Frequent training and interaction with the workers to make them conscious about
their work and different safety procedures.
The commissioning of a modern chemical plant is a complex and difficult exercise, the final
stage of a major project involving the authorization, design, construction and start-up of the
plant. Commissioning is the time when the quality of work carried out during earlier phases of
the project is validated, and also to confirm a return on the investment commences.
All chemical plant projects pass through a similar sequence of phases between the original
conception of the project and the beneficial production of chemicals. Decisions taken throughout
the whole of the project as per specification are complex and are based on the best information
available at that time. This information is sometimes incomplete or conflicting due to uniqueness
of concept to commissioning. The correctness of these decisions is not apparent until the plant is
commissioned and is into beneficial operation.
The detailed design stage (operation and engineering) is the one at which the hardware of the
plant is constructed and is defined in detail. The whole of the design clearly has a major
influence on the success of commissioning but there are some features that have a large effect on
commissioning which were overlooked.
P a g e 10 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Major Challenges overcome during Pre-commissioning / Commissioning and
start up phase of the project:
 Readiness of Utilities is a key factor to start the Pre-commissioning/Commissioning but the
whole scenario was very different for BCPL due to lackluster approach of LSTK contractors.





Storage limitation of Raw water reservoir and high silica content in raw water intake
well was major reason for delay in commissioning of DM plant and CPP – Initial
use of stored rain water from Raw water reservoir in DM plant to supply desired
specification of DM water to CPP to overcome the delay in commissioning of CPP
and Steam generation which in turn help in supplying Power and required steam for
steam blowing of VHP/HP/MP/LP steam lines. Further OBR (Output between
regeneration) reduced and regeneration cycle increased in DM plant.
Scarcity of Nitrogen due to delay in commissioning of Nitrogen plant. Liquid
Nitrogen/Nitrogen cylinders were procured from other agencies to mitigate the
demand for Passivation, Inertisation, Seal gas, blanketing etc.
Card board blasting of different pipelines were carried out with the help of portable
air compressor hired from other agencies due to non-availability of plant/instrument
air.
Hydro-blasting/ Rotomould cleaning/Chemical cleaning and degreasing were
carried out to clean the pipelines for different services.
Passivation of cooling water circuit got delayed due to non-availability of heat load
in the circuit.
 Improper Storage of different equipments, Pipes/fittings, structural steels etc. resulting in
equipment damage. Commissioning of Refrigeration package was delayed as precommissioning activities required more time than anticipated as the equipment supplied by
vendor/mechanical contractor was not properly protected and cleaned.




Breakdown of Utility Boiler-2 in CPP during commissioning.
Damage of 68 Pole structures of the HT Overhead transmission line during cyclone.
Leakage in cold box of ECU during commissioning.
Delay in commissioning of other units i.e. GDU & GSU, Lakwa unit due to remote location.
 Different Feed Stock Issues:
 Shortage in supply of Natural Gas and polymer grade Naphtha - feed stock to
Ethylene Cracker Unit.
P a g e 11 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.



Continuous follow up done with supplier end for the natural gas. M/s.NRL supply
naphtha by procuring it through different external resources. More over the
additional shortage is overcome by procuring it directly from different agencies.
Shortage of Butene-1 - feed stock in downstream PE Unit, which limits the
production of a desired grade of polymer.
To overcome this issue deliberations made with different industries to procure the
same as per specification. Presently it is being received from various domestic
sources. Further BCPL is looking to line up butene-1 through bulk imports and
also planning to set up a 10000 MT Butene-1 plant.
Shortage of Propylene feed for Poly propylene Unit because oflow in house
production of propylene due to scarcity of feed stocks i.e. naphtha and natural gas
with desired composition. This shortage is overcome by procuring the same from
other sources.
Inadequate availability of required specification Pentane leading to procurement
from limited parties.
 Problem in start-up of Polypropylene Unit due to problem in inlet dryer. Modification done
in the Dryer outlet to recirculate maximum amount of propylene to propylene storage sphere
during initial start-up of PPU reactor.
Plastic Processing Sector in the NE Region- a short analysis
Considering the emerging scenario on the industrial front, as a result of the industrial policy
being pursued and other related developments, the following SWOT analysis has been cast for
the plastic processing units in the region:
Strengths:
 Competitive feedstock sourcing is now available with continuous operation of Brahmaputra
Cracker & Polymer Limited.
 The North Eastern region is now considered as a corridor for exports of plastic/polymer
products through cross-country trade with the neighboring countries like Myanmar and
Bangladesh.
 Favorable investment decisions from some well-known players in the plastic/polymer field
already in existence.
 No desulphurization unit is required as North-East crude does not contain any Sulphur.
Weaknesses:
 The size of local market is rather small and spread out. There is a general tendency
(especially for major investors) towards concentration in one or two states.
P a g e 12 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
 The region represents one of the least industrialized zones, thus restricts the range and size of
market for products catering to industrial use. Universally, the industry sector accounts for
the major share of plastic use.
 At present, access to the tooling, mold making and other infrastructural facilities and plastic
processing machinery manufacturers is difficult.
Opportunities:
 Opportunity to establish cluster units/plastic industrial estates, or ‘poly parks’ with inherent
overall operational advantages.
 Opportunity to replace the existing expensive, energy intensive and/or depleting products
such as steel, timber, etc., by plastics products
 State electricity board got opportunity to get electric power from BCPL Captive Power Plant
through grid synchronization to overcome the additional domestic power requirement.
 It lays the basis for the development of a new industrial logistics and transportation system in
the North-East, and make further petrochemical plant additions /expansions much cheaper
and easier.
 Drive employment in a more sustainable manner and it will train a new generation of
workers, create opportunities for Entrepreneurs, and drive research dollars in the
neighborhood.
 Different agencies are coordinating with BCPL to set-up a CO2 recovery system to utilize the
CO2 emitting from upstream gas processing unit of BCPL for blanketing the Crude oil well.
Threats:
 Competition from imported products.
 The feedstock sourcing scenario will get changed for the better, with the continuous
operation of Brahmaputra Cracker & Polymer Limited complex. It can indeed make the
region a very competitive sourcing point for a range of polymer feed stocks.
Logistics- A case study
Before detailing on the challenges and innovative solutions being taken into consideration, the
detailed discussion on following topics are required.
a. Supply chain vision and scenario:
“To emerge as a dominant petrochemical player in the northeast region, providing value to
stakeholders, offering best-in class products & services, contributing to economic growth while
remaining environmentally conscious”.
P a g e 13 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Achieving supply chain collaboration through:



Least human intervention in demand planning using tools and mechanism such as
predictive advance forecasting methods, suppliers and logistics provider database and
statistical forecasting methods.
Forward looking supply management using control tower method for real time
synchronization with demand, anticipating what-if-scenarios for quality and supply
failures and predictive maintenance process across the supply network.
Ensuring highest quality at lower cost by proactive collaboration across all supply chain
partners to leverage scale and segmentation of logistics to deliver as per the requirements
on real time.
We procure various feed-stock materials on regular intervals as per manufacturing planning
from various sources including manufacturers across domestic and international locations
through competitive bidding as per CPSEs procurement guidelines. Basic purpose for getting
regular sourcing is defined as per below conditions:




Chemicals have definite life span (Shelf life) resulting in expiry if stored for longer
storing period and its preservations is not technically reasonable.
Costs of sourcing is considered to be high because of related costs of logistics and
storage.
Limited availability in both international and domestic markets.
Dynamic nature of consumption pattern.
Figure: Chemicals and catalysts procurement
b. First step filtering:
There are many chemicals and additives required depending on the production capacity and
product variants. Amongst the long list of chemicals few of the critical feed-stock chemicals
primarily required for plant operations areNaphtha, Butene-1 and Propylene. For the purpose of
P a g e 14 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
this case study, deciding on the most critical feed-stock chemical is made on various factors as
follow:






Production capacity of Polyethylene (both HDPE & LLDPE) is 226,000 Tonnes per
annum which is more as compared to 60,000 Tonnes per annum of Polypropylene
(PP).
Advance technology licensed by INEOS Technology capable of producing various
PE grades having MI value in the range of 0.05 to 21 g/10 min without any technical
problems.
Scarce availability of feed-stock chemicals for PE in both domestic and international
market as per the required specifications.
Chemical feed-stock have wider tolerances on technical specifications making it
convenient to source from manufacturers operating under various licenses.
Consistent and high demand of film grade PE in domestic market.
Better earning with profit margin on PE grades sales.
Considering the above factors as the first level of filtering, it is found that for PE plant Butene-1
is the critical feed-stock chemicals required for producing various grades of highly demanding
polymer grades.
c. Selection of critical feed-stock chemical:
Butene-1 is required as co-monomer in LLDPE/ HDPE plants to maintain the density and quality
of polyethylene polymers. Annual requirement is 8900 MT as per the BDEP specified by the
Licenser. Target level for producing PE grades polymers, requirement of Butene-1 is required on
an average of 50 to 60 MT per day for production of highest selling grades i.e., Film grade.
Due to untimely and interrupted supplies of the chemical, plant production is getting affected on
regular basis thereby leading to huge operation losses amounting to crores of rupees. The major
concern is absence of own Butene-1 producing plant and dependency on external supply sources
from both domestic and international markets.
Considering the crisis situations of the plant operations Butene-1 is sourced from domestic
markets mostly from major petrochemical industries. For maintaining consistent supplies an
agreement was made with its major promoter for 5,000 MT quantity to partially fulfill the annual
requirement at 8,900 MT.
d. Problem statement:


Due to technical issues at promoter’s plant, supplies got disrupted affecting the plant
for a considerable period of time.
On emerging situations, various sources were being explored based on competitive
bidding process. Based on the tender, limited manufacturer participated and the
intended purpose didn’t meet.
P a g e 15 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.






One of the major Indian petrochemical industries could not maintain the continuity of
supplies considering technical issues and finally became the dormant supplier.
Alternatively, another major Indian petrochemical industry could not maintain
supplies as they do not have their own production facilities and is totally dependent
on their import consignments.
Supplies to BCPL even after consistent pursuit, never happened to be regular due to
what-if-scenarios could not be made as there were substantial and unprecedented
delays in availability from sources leading to improper logistics arrangements.
Inadequate in-house storage capacity.
Requirement of specialized nature of tankers with statutory clearances.
Poor weather conditions, incessant rainfalls and flood conditions.
e. Decision modeling of sourcing and logistics:
Before making decision on ‘how-to-procure’ Butene-1 for maintaining fulfillment with cost
effective and responsive supply network, supply chain mapping was being considered keeping
infollowing major points. Further to that focus on ‘how to use’ our unique position in the entire
supply chain network to build a collaborative logistics programme, driving deeper collaboration
between our key partners to tackle the critical sustainability issues to tackle empty tankers,
wasted kilometers and avoidable CO2 emissions.
i. Multiple sources enables continuity of availability.
ii. Increase in supplier base to reduce dependencies on a single source of supply.
iii. Actualization on availability of minimum quantity on daily basis at supplier end.
iv. To achieve competitive sourcing costs.
v. Optimization of logistics costs
vi. Fast fulfillment and supply network.
vii. Reduction in frequent supply network re-design.
f. Analysis and executed steps:
The purpose of this paper is to analyze and devise solutions what problems and challenges BCPL
is facing in the implementation of a global sourcing strategy for feed-stock chemicals viz.,
Butene-1 to sustain the plant operations.
Due to the global nature of activities in sourcing and logistics, which in most cases are continents
apart, presents a huge challenge. The modes of transportation normally used are mainly through
high seas and inland roadways or rail. In most of the instances, multi-modular shipment have
been utilized before reaching the final destination. The long distances between parties coupled
with slow methods of transport increases shipping and inventory carrying costs. Such challenges
necessitate that workable methods needs to be devise through innovative ways to meet the
highest levels of service and be cost-effective. A true solution should recognize the broad range
of entities with views as it a whole rather than individual links to enable efficient and profitable
operations across the entire business.
P a g e 16 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
Materials team began by conducting a detailed assessment of the supply network within each
strategic business unit keeping in mind the constraints. The goal was to create a step change in
how the company operated across six areas:





Planning and operating processes
Supply chain network and logistics
Organizational structure
IT enablement
Performance measurement
Based on the challenges faced by the plant in sourcing and logistics, following steps are taken for
administration and execution to reduce the uncertainty in demand and supplies.








Enabling the entire procurement procedure through IT enabled software viz., SAP
system.
Implementation of e-tendering process for procurement of supplies and services.
Engaging MSE vendors of the region for maximum participation in tendering activities.
Implementation of All India rate contracts for hiring of transportation services.
Standardization of tendering process as per the laid guidelines.
Implementation of Procure to pay (P2P) process.
Introduction of online payment mechanism.
Involvement of cross functional teams for monitoring and suggestion purpose related to
sourcing and logistics.
a. IT enabled processes:
Technology has transformed production-led supply chains to consumer-centric demand chains.
Corporate need to better integrate requirements and supply chain decision-making to optimize
the distribution network. Processes are aligned to achieve the key transformations like:




End to end lead times.
Supply base management.
Network optimization and transportation.
Time to market.
Activities executed through IT enabled environment using SAP R/3 platforms are:






End to end sourcing planning with vendor performance appraisal.
Procurement activities by e-tendering through various tendering activities viz., open
domestic/ international, limited domestic, board purchase, petty purchase etc.
Day to day basis consumption pattern analysis.
Sourcing lead time evaluation and strategy development for order grouping.
Inventory analysis and order fulfillment status management.
Quality management and performance analysis.
P a g e 17 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
In BCPL, access to low-cost raw materials (or feedstock) is an important competitive advantage
for which extensive study and survey was made. As the industry is facing volatile feed-stock
costs, cyclical product prices and increasing pressure to reduce its impact on the environment,
optimization of end to end supply chain was critical to achieve operational excellence and the
overall objective of maximizing return on capital while ensuring safety and sustainability.
Through planned and systematic usage through IT enabled processes and methods, the objectives
as stated above was achieved with reduction of order to deliver cycles by 15% as compared to
preceding year.
Enabling IT in the entire sourcing practices have resulted in to 80% delivery with track and
trace, 15 freight forwarders connected, 85% orders placed digitally, purchasing analysis made
simpler, extracting best in supply chain optimization with digital end to end planning, forecasting
and inventory management.
b. Logistics arrangements for supply chain efficiency:
In Sales & Operation planning, we had evaluated various options in getting detailed analysis on
what to produce and where, and what feedstock to purchase and how to transport and store them.
It is imperative that intelligent planning of an integrated petrochemical complex is essential for
operational excellence and customer satisfaction. Yet, true optimization is beyond the capacity of
traditional planning systems, when plans must take into account the cost of alternative feedstock
and many production constraints. We have appliedvarious tools which is configured to model the
capacity, processes and constraints unique to our business in order to provide truly optimal
production schedules.
P a g e 18 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
However, for running Film Grade consistently at approx. 70% load, Butene-1 consumption is
expected to be around 1,500MT per Month and with 90% load, the requirement would be
1,900MT per month. Since the market demand is more for film grade and HDPE grades are not
in much demand, the consumption of Butene-1 is on the higher side and may vary accordingly.
Considering above precarious situation and to meet the urgent requirement of Plant, following
key activities are performed.



Network design containing shipment mode and storage facilities analysis.
Keeping two way communication channels open with suppliers and service providers.
Facilitating synchronization of requirements and spend analysis with internal
departments.
Moreover, sourcing from international markets was found to be costlier as compared to domestic
sourcing due to heavy detention charges (in case there is delay in logistic operation, hiring
storage facility at port, customs and Port clearance, etc.) and inventory holding of 30-40 numbers
of specialized shipping containers in circuit/ loop to meet the minimum monthly target of 1000
MT. Reduction in shipment lead time is also not possible which adds to the cost besides
demographic disadvantages of the plant (i.e. North-East region).
Calculations on transit time and logistics issues in Import shipments
I.
Weekly network design and evaluation activities carried out on excel as
preliminary exercise to manage the sudden disruptions at sources and service
providers engaged.
The network modeling is made to optimize sourcing foot-print, transportation and lead times. It
has ensured full visibility and ensure timely order fulfillment.
P a g e 19 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
II.
Daily MIS on in plant production capacity and in-plant consumption daily
basis.
III.
Daily MIS reports on loading & unloading and receiving status.
P a g e 20 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
IV.
Daily in-plant receipt with stock status
V.
Daily cumulative receipt status from various sources of supply.
P a g e 21 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
A detailed logistics network and inventory model designed allowed to evaluate different
scenarios that would minimize intra-unit freight costs and optimize inventory levels at each
stocking point. The model also supported our redesign of the company's port facilities to improve
efficiency and accommodate its growth plans. After assessing the value of routes and hubs, we
recommended that the company selectively expand existing sources into hubs and add additional
hubs in high-value areas.
Conclusion:
Brahmaputra Cracker & Polymer Limited will prove to be the potential game-changer in shifting
the focus of plastic industry to North-Eastern states of India and become the hub for future
growth. BCPL will ensure consistent local supply of raw materials mainly, Polyethylene and
Polypropylene. This local supply of raw material would enable the competitiveness of local
plastic processors thereby stimulating growth of downstream units in the region. We believe this
project is a catalyst for future use and development of significant volumes of Natural gas
available in Assam.
The just commissioned Brahmaputra Cracker & Polymer Limited (BCPL), which produce more
than 1 lakh tonnes of Polymer in the first year of operation of which more than 50 % are prime
grades is ready to supply the major feed stocks for the polymer industry in full flow, and started
showing its favorable impact on the plastics/polymer industry in the region, with around 10-15
medium scale units are coming up.
One of the first insights was the degree to which a tightly integrated planning discipline can drive
reliability and customer value. It was recommended dynamic scheduling rather than the current
monthly recalibration, which was preventing the company from quickly addressing changing
customer requirements.
Due to innovative steps taken for sourcing and logistics operations, availability of Butene-1,
Naphtha and Propylene are maintained at the minimum required quantity on monthly basis
without any major disruptions in supply chain. This has resulted in meeting monthly production
targets as set in S&OP.Although the discussed issues are being mitigated through enterprise
logistics solutions, still they do have sustainable effect on supply chain performance of
sourcing.Indian petrochemical industry has unrealized potential. Polymer demand is expected to
P a g e 22 |23
21st Refinery Technical Meet
Strategy in setting up a Petrochemical Industry in NE India and its
journey towards Global Competitiveness: A study.
grow by 7% in 2014-15 with a healthy growth in the relevant industries such as clothing,
automobiles etc. Government and the industry players will have to work in tandem to achieve
ambitious targets for the industry.
These are truly exciting times in the North-Eastern States! A manufacturing renaissance has been
launched as a result of this opportunity.
Disclaimer:
(The information upon which this paper is based comes from our own experience, knowledge
and data bases, supplemented by reference to primary sources and published industry data. Any
opinions expressed are those of the author/co-author as of this date. They have been arrived at
following careful consideration and enquiry but we do not guarantee their fairness,
completeness or accuracy. We do not accept any liability for your reliance upon them).
References:
1. Report of the Sub-group on Petrochemicals for the 12th Five Year Plan.
2. India Petrochemicals Industry Outlook to 2015.
3. Handbook on Indian Chemical Industry, IndiaChem2019 and 2012.
4. www.cipet.gov.in.
5. IndiaChem Gujarat 2012.
6. Crisil Research.
7. Chemicals& Petrochemicals statistics at a glance: 2013, GoI.
8. Report on Indian Plastic Industry 2012 - 2013, Plastindia Foundation-2015 report.
9. Asia petrochemical industry conference-2015 report.
P a g e 23 |23
Energy Efficiency and Reliability Improvements in Design of
Grass-root CDU/VDU at HPCL Visakh Refinery
Deepak Kumar Jha
Ashraf Jamal
Hindustan Petroleum Corporation Limited, HQO, Mumbai
1. Preamble:
HPCL presently operates an 8.33 Million Metric Tonnes Per Annum (MMTPA) refinery at
Vishakhapatnam, Andhra Pradesh, India. It is intended to further enhance its refining
capacity to 15 MMTPA under Visakh Refinery Modernization Project (VRMP).
The basic process configuration envisaged under the VRMP is a new Crude Distillation &
Vacuum Distillation Unit as Primary Processing unit, Full Conversion Hydrocracker as
Secondary Processing unit, Residue Upgradation Facility as Bottoms Upgrade unit and a
new Naphtha Isomerization unit along with revamp of existing units for entire BS-VI
specification fuels production.
The major energy consuming processes in refineries, in order of overall energy consumption
are typically CDU/VDU, Hydrotreaters, Reformer, Catalytic Crackers and Hydrocrackers etc.
Out of which CDU/VDUs are considered to be the largest energy consumers, utilising around
20% of a refinery’s total energy consumption, depending on configuration and type of crude
processed.
Competitive energy efficiency benchmarking data indicate that there is a potential of
economic improvement in energy efficiency of the erstwhile refineries. These potential
improvements will amount to huge annual costs savings and improved energy efficiency will
also result in co-benefits as absolute reduction in emissions.
M/s Engineers India Limited was appointed as consultant for design of new grass-root 9.0
MMTPA CDU/VDU.This paper describes details of various energy efficiency and reliability
improvements considered in design of Grass-root CDU/VDU at HPCL Visakh Refinery.
Page 1 of 9
2. Brief Description:
Energy efficiency and reliability has become an essential feature in the design of process
plants due to increasing cost of energy and stringent environmental regulations. During the
design basis finalization of the new CDU / VDU of 9.0 MMTPA, various energy efficiency and
reliability improvements were targeted viz., energy efficiency consideration during design
crude selection, two stage desalter, process integration for higher preheat temperature (coil
inlet temperature, CIT), air-preheating, distillation optimization, liquid ring vacuum pump,
vacuum type selection etc.
Following major aspects, considered for improving energy efficiency and reliability of the
CDU/VDU, are discussed in detail:
 Preheat train is an integral part of CDU/VDU and considerable heat is recovered
through preheat train. The preheat temperature of 300-305 deg C is considered as
very good heat recovery; however, in HPCL CDU/VDU by optimizing and improving
the integrated heat exchanger network of crude oil Preheating Process, preheat
temperature of 320 deg C is targeted and achieved with optimal preheat integration.
 Consideration of Organic Rankine Cycle (ORC) module for low level heat recovery in
order to improve overall energy efficiency of the CDU/VDU. This measure will
generate 2.4 MW power from waste heat which was getting dissipated through air
cooling or water cooling.
 Economic analysis for Liquid Ring Vacuum Pumps (LRVP) in Hybrid ejector system
(two stage eject and third stage LRVP) vs. three stage steam ejector system and four
stage steam ejector system. With overall reduction in Opex, Hybrid ejector system
has been considered in the design. Based on preliminary economic analysis, this
measure has resulted in substantial reduction to the tune of 8 TPH in steam
consumption in CDU/VDU with IRR of more than 20%.
 Fouling Management: Consideration of Helix exchanger in dirty/viscous services for
better reliability and run length of the exchangers for heavy services as fouling is one
of major cause of energy in-efficiency in the crude preheat trains. Antifoulant injection
at strategic location in crude and vac residue circuit has been considered for
improving the reliability of heat exchangers in fouling circuits.
Page 2 of 9
3. Detailed Approach for improving Energy Efficiency and Reliability of CDU/VDU:
Energy is one of the major Opex contributor of any refinery. An overall integrated
optimization approach has been adopted during the design stage of VRMP where in capital
expenses can be made to reduce energy consumption resulting in overall benefits in project
life cycle cost. It is well known that designing of energy efficient process unit during grassroot phase is more capital efficient than attempting to retrofit existing plants.
In view of the above, HPCL has targeted energy efficiency and reliability as cornerstone of
mega projects such as VRMP with a major focus area to reduce energy use. In a refinery
project, typically improvement ideas are either not identified or not implemented due to
concern of additional project costs or time implication in project implementation. To produce
an energy efficient design, energy optimization must be envisaged during the conceptual
stage of the project itself when it is possible to make the biggest improvements to the
efficiency of the design at marginal cost. These unexploited opportunities of energy
improvements due to conventional approach of project implementation results in munching
the profit margin in future and endangering the sustainability of the business.
3.1 Design Crude Selection:
Extensive brainstorming exercise has been carried out for selection of the design crude for
CDU/VDU. Crude basket for the Visakh Refinery, having approximately 130 crudes, is
screened for selecting the base crude for CDU/VDU design. Risks with opportunity crude
handling is also strategically assessed along with targeting the mitigation of the challenges
associated with processing these opportunity crudes in order to maximize gross refinery
margin (GRM) for business sustainability & profitability.
Opportunity Crudes are typically available at discounted prices due to its intrinsic processing
issues viz., high acidity, high sulphur content, nitrogen, aromatics and low API gravity etc.
HPCL has set target to consider state of the art upgradation processes to process these
crudes in energy efficient, cost effective manner. It is imperative that higher capital
investments will be required for processing opportunity crudes; however, the economic
benefits offered by these opportunity crude oils are considerably high. Typically, any Crude
Distillation and Vacuum distillation unit is designed of TAN 0.5, however, upgraded
metallurgy to handle high acid crude up to TAN 1.0 has been considered in designing of the
Page 3 of 9
CDU/VDU. Upgraded metallurgy considered is 3.0 MM SS316L in bottom section and 3.0
MM Monel clad in top section of Crude column. For Vacuum column, shell is considered as
KCS cladded with SS316L including top and bottom head. All internals are SS316L with Mo
more than or equal to 2.5 wt % including trays and packing bed in Vacuum Column.
Upgraded metallurgy is also considered for other pieces of equipment having potential of
naphthenic acid corrosion.
Base staple crudes have been shortlisted as crudes in the API range of 27 to 30 from the
crude basket of Visakh Refinery considering various factors viz., API range, Sulphur, Vac
resid generation, historical prices, future availability, and geographic location. Basrah Light
and Arab Heavy were shortlisted as the two base crude. Further, few opportunity crudes
(heavy Crudes) have been shortlisted based on API less than 27. The criteria for selection
of crude mixes was:

Vac Residue yield for the selected crude mix should be in range of 25% ‐ 30% (in
view of downstream bottom upgradation capacity handling)

Total sulfur content of the selected crude mix < 2.9 wt% to satisfy sulfur and
hydrogen balance
Basrah Light & Doba (Ratio 87: 13) and Arab Heavy & Dar Blend (Ratio 92.5: 7.5) has been
arrived as design crude case for the VRMP CDU/VDU design.
3.2 Crude Preheat Train Temperature Maximization:
To optimise waste heat recovery and efficient heat integration, Pinch Technology is used to
analyse heat availability in process’ hot streams and to match it against the heat demand of
suitable cold streams, in an optimum fashion. This optimizes the preheating of the cold
streams by using hot streams’ waste heat, and saves fuel in furnaces.
A maximum preheat temperature of 305 deg C could be achieved in few of the latest
CDU/VDU design in last decade in India with design crude in API range of 25-30. The
preheat of 300-305 deg C is considered as very good heat recovery in such designs;
however, in HPCL CDU/VDU by optimizing and improving the integrated heat exchanger
network of crude oil preheating Process, preheat of 320 deg C was targeted.
Page 4 of 9
Reference case of CDU/VDU design was considered as one of the latest 9.00 MMTPA
CDU/VDU design carried out in the recent past for HPCL. In this design, a maximum preheat
temperature of 303 deg C was achieved for Crude of API 30. The total preheat circuit was
divided in three preheat train network i.e. preheat train -I from crude charge to Desalter,
preheat train-II from Desalter to preflash drum and preheat train -III from preflash drum to
Crude Heaters. The total duty imparted in preheat network was approximately 176
MMKcal/hr i.e. 58 MMKcal/hr, 23 MMKcal/hr, 95 MMKcal/hr in preheat train I, preheat train II
and preheat train III respectively. Accordingly, the Crude Heater and Vacuum Heater duty
was arrived as 91 MMKcal/hr & 46 MMKcal/hr.
For maximization of preheat temperature to meet the target of 320 deg C in VRMP
CDU/VDU, EIL was advised to carry out extensive pinch analysis with various options of the
crude preheat network to improve energy efficiency.
In this VRMP CDU/VDU design, the total preheat circuit was divided in two preheat train
network (omitting the Costly Preflash Drum as heavier design crude are considered for
VRMP CDU/VDU design) i.e. first preheat train from crude charge to Desalter, second
preheat train from Desalter to Crude Heaters. The total duty imparted in preheat network
was approximately 191 MMKcal/hr i.e. 57 MMKcal/hr in preheat train I and 134 MMKcal/hr
preheat train II respectively. This has resulted in preheat temperature of 320 deg C. With this
preheat maximization, Crude Heater and Vacuum Heater duty was reduced to 82 MMKcal/hr
& 42 MMKcal/hr respectively without any impact on product yield & specifications. However,
to achieve this target, approximately eight new Heat exchangers were introduced in preheat
network (than previous design) to maximize the potential of heat integration; nevertheless,
this has resulted in reduction of overall Crude heater and Vacuum heater duty by 10%
resulting in huge fuel savings. Preflash drum has been kept as future revamp requirement.
Utilization of crude column overhead vapor in preheat in the first leg of the preheat train with
resulted overall benefit in preheat. Erstwhile, the vapours from crude column overhead are
partially condensed using air coolers wherein the heat is lost to the atmosphere. However,
wwith increased energy cost, recovering heat from Crude Column Overhead Vapor would be
very much economical. Providing this crude column ovhd exchanger in preheat has
recovered approximately 20 MMKcal/hr of duty. Though, this provision was considered in
both CDU/VDU BDEP designs.
Page 5 of 9
Table: Comparative Analysis of Reference CDU/VDU Design Case & VRMP Case
Sr.
No.
Case
Heat Recovered in Preheat trains
Preheat
Train-I,
MMKcal/hr
Preheat
Train-II,
MMKcal/hr
Preheat
Train-III,
MMKcal/hr
Total Heat
Gain in
Preheat,
MMKcal/hr
Preheat
Temp.
Deg C
Crude
Heater
Duty,
MMKcal/hr
Vac Heater
Duty,
MMKcal/hr
1
Reference
Case
58
23
95
176
303
91
46
2
VRMP
Case
57
134
--
191
320
82
42
As it is intended to process the opportunity crude, it was targeted to have flexibility to
operate the desalter at an optimum temperature to maintain peak desalter performance. To
avoid asphaltene precipitation associated with opportunity crude processing, the preheat
train’s is designed to have flexibility of to change the desalter temperature by 15-20°C.
3.3 Low Level Heat Recovery (Organic Ranking Cycle Implementation):
Considerable amount of waste energy lies in a low temperature range which is dissipated
through air cooling or/and water cooling. With the increased energy costs, it is felt prudent to
recover low level energy lost to the atmosphere to increase overall energy efficiency of the
CDU/VDU. Thus, Organic Rankine Cycle (ORC) module for low level heat recovery has
been envisaged.
ORC is a technology that operates similarly as the Steam Rankine cycle, except that the
ORC uses an organic working fluid instead of steam. This organic working fluid has a lower
boiling point and a higher vapour pressure than water, which is suitable to use low
temperature heat sources to run a turbine for power generation. The Organic Rankine Cycle
makes use of and organic fluid with a boiling point lower than water. The fluid enables
recovery of heat from lower temperature sources. The low temperature heat is used to drive
a turbine and create electricity.
In CDU/VDU design, a screening analysis has been carried out for identification of all the
potential streams; there are few low level hot streams which are getting cooled in Air Fin
Coolers (AFC) and Water Coolers after preheat maximization. Heat of these streams cannot
be recovered in regular manner. Hence, HPCL appointed a consultant to carryout feasibility
to assess the potential of using Organic Rankine Cycle (ORC) technology to recover low
grade waste heat of these hot streams and generate electricity.
Page 6 of 9
It has been observed that approximately 20 MMKcal/hr of low level heat is available in
temperature range of 170 deg C thru various CDU/VDU product streams i.e. Vac Diesel,
HVGO, ATF/Kero, HGO & VR+Slop. It has been estimated that approximately 2.4 MW
power can be generated. The ORC module will result a simple payback period of 4 years @
power price as low as Rs. 6.5 Rs/kW. It is further intended to develop BEDP for this facilities
and implementation in CDU/VDU. Area requirement for ORC module were already earmarked during CDU/VDU BEDP preparation.
3.4 Energy Efficient Hybrid Ejector System
In CDU/VDU unit, vacuum in VDU column is typically achieved by use of steam ejectors.
Steam ejectors are though reliable but inefficient, and these ejectors are typically considered
as vacuum generation machine due to easily available steam in refineries. With increasing
steam price, steam ejectors are being replaced by Liquid Ring Vacuum Pumps (LRVP).
LRVP is a constant volume machine having isothermal compression commonly uses water
as a seal liquid since it can be separated and reused safely.
Steam jet ejectors are mass flow devices for vacuum services with no moving parts
operating on the principle of momentum. Ejectors are reliable if the process gas and motive
Page 7 of 9
steam conditions are constant. Any changes in the process gas composition or flow rate
affects the steam ejectors performance and stability.
Hybrid systems can be considered for the applications in which deep vacuum is required and
system reliability is critical. Integration of a steam jet ejector’s high vacuum capability with a
liquid ring pump’s stability can provide a system with stable operation during process upsets,
increased reliability and lower total operating costs.
LRVP are generally more expensive compared to steam ejectors; however, for better
operating cost savings, LVRP is typically considered in Hybrid ejector system as
replacement of the last one or two stages of a multistage steam ejector system. Under
VRMP CDU/VDU, various options for developing energy-efficient vacuum generation
systems were studied viz., 3 stage steam ejector system, 4 stage steam ejector system and
Hybrid ejector system (two stage steam ejector followed by 3rd stage as LRVP). It has been
found that hybrid ejector system has resulted in good amount of reduction in steam
consumption. Thus, a Hybrid ejector system has been considered in VRMP CDU/VDU to
meet Vacuum Column Flash zone pressure at 39 mmHgAbs.
3.5 Fouling Management:
Fouling is an important factor for efficiency losses in the CDU/VDU and the crude preheat
train is prominently susceptible to fouling. It is also worth noting that time based increased
fouling has severe implication of preheat temperature that has resulted in increased firing in
the heaters resulting Opex increase and poor energy efficiency. Reducing this additional
heating load could result in significant energy savings. One of the commercially available
heat exchanger designs that have a lower fouling rate compared to the traditional shell &
tube heat exchanger is Helix exchanger. These exchangers generate the Helical flow with
increased turbulence. Overlapping baffles define a helical path. Flow is divided & rotated,
like a static mixer. Low fouling design aims at maintaining the design heat duty resulting in
significant run length between cleaning.
Helix Exchangers have been considered in all preheat exchangers in Vacuum Residue
service. Moreover, antifoulant injection is also considered into Crude line at the common
discharge of Crude Charge pump and into VR+Quench line at Common suction of
VR+Quench Pump for improving the reliability of heat exchangers in vacuum residue.
Page 8 of 9
Heavier crudes generally contain more salts, making desalting more important in refineries.
This salt needs to be removed so that the crude oil can be processed in a refinery without
fouling heaters and exchangers and other equipment. Poor desalter performance is a
common problem that may be encountered while processing extra-heavy crudes. This may
lead to severe crude column overhead corrosion, including exchangers, piping and vessels.
The desalter in VRMP CDU/VDU is designed to handle feed salt content of 57 Ptb and feed
BS&W of 2.0 vol.% (max). Preheated crude enters the two stage desalting system under
desalter pressure control. The salts are then dissolved in the water by mixing of preheated
wash water in a mixing valve and an electric current is used to separate the water from the
oil. Sufficient pressure is maintained at the desalter to ensure no vaporization. Desalting will
reduce corrosion and minimize fouling of process units and heat exchangers. The efficiency
of desalting is influenced by the pH, gravity, viscosity, and salt content of the crude oil, and
by the volume of water used in the process.
4.0 Summary & Conclusion:
In an increasingly competitive global refining business environment, it is the need of the hour
to reduce the operational costs by evaluating & scouting opportunities for energy savings at
the design stage of Project itself. Cost-effective investments for energy efficiency
technologies and practices will meet the challenge of both maintaining the high quality of
products while reducing production costs.
In Grass-root 9.0 MMTPA CDU/VDU at HPCL Visakh Refinery, Energy Efficiency and
Reliability Improvement targets are achieved by various above mentioned considerations
which has resulted in the overall specific energy consumption reduction by 10%.
It can be summarized that substantial energy savings can be achieved in grassroots plants
and much of the improvements exist in the process integration, energy efficiency
improvement & reliability improvement of CDU/VDU. The key to achieving this is to adopt
new approach and methodology of optimisation in design phase of the project in order to get
the energy improvements.
****
Page 9 of 9
PAPER FOR 21st REFINERY TECHNOLOGY MEET (RTM), 2017
Paper Title: “Increasing Hydrogen Recovery in PSA through Adsorbent Replacement:
A First-of-its-kind Effort by HPCL”
Authors:
 Arun Kuniyil, Prashant Mishra, T A Rajiv Kumar, S N Sheshachala, Dr. Peddy V
C Rao, Dr. N V Choudary, G Sriganesh – HPCL Green R&D Centre, Bengaluru
 S Tazeem Abbas, Ashok P Golekar, S V Choulkar, Mandar A Joshi – HPCL,
Mumbai Refinery
ABSTRACT:
Hydrogen is one of the critical process utility in refineries used in processes such as Diesel /
Naphtha Hydrotreating, Hydrocracking, Isomerization and Lube upgradation. With increasing
stringency on fuels specifications and with increasing bottoms upgradation processes,
Hydrogen consumption has significantly increased in refineries making it a high value utility.
Pressure Swing Adsorption (PSA) is the widely used technology for H2 purification in
refineries. The technology is licensed by very few multinational licensers and all its critical
details are closely guarded.
In view of the rising demand for Hydrogen in refineries, a need for an indigenous PSA
technology was felt and HPCL R&D along with its Korean collaborator has developed PSA
technology for H2 purification from refinery gases. HPCL’s commercialization efforts
emphasize that adsorbents selection and cyclic sequencing are highly critical for a sustained
PSA operation.
As a step to reinforce its capabilities in PSA technology, in a first-of-its kind effort, a project
was taken up by HPCL for selection and replacement of adsorbents at one of its existing
commercial PSA unit in Mumbai Refinery. The PSA unit was designed to process 33400
Nm3/hr of feed gas coming from the steam reformer of HGU and to produce 99.99 mole%
Hydrogen. The PSA was in service since its commissioning in 2000 and the adsorbents
have not been replaced since then. However, the Hydrogen recovery of the unit has
progressively deteriorated due to ageing of adsorbents. HPCL R&D having acquired
expertise in PSA technology including adsorbents, has carried out research studies on
numerous variety of adsorbents and selected the optimum adsorbents. These were loaded
in the PSA unit during the T&I. The unit was restarted successfully and is presently operating
meeting the design product purity and recovery. This paper shares HPCL’s experience in
this unique effort of adsorbent replacement and achieving successful performance.
Introduction
Hydrogen is one of the premium utility used in refineries and is required for major
upgradation facilities such as Diesel Hydrotreating, Naphtha Hydrotreating, Hydrocracking,
Isomerization and Lube upgradation. Over the past years, consumption of Hydrogen has
substantially increased in refineries. The main reasons can be attributed to the requirement
of processing heavier crudes and bottom of the barrel up gradation for obtaining better
GRM’s, as well as for meeting the stringent fuel specifications in regards to the growing
environmental concerns.
Pressure swing adsorption (PSA) process is a well-recognized process / technique used for
separation of gaseous mixtures and is well employed in refineries for purification of hydrogen
from mixture of gases. It is widely known that the adsorbent material employed in a PSA
process is extremely important in defining its properties, and process engineering for precise
control is required for enhanced performance of the process. Hence the technological
intricacies of this process, expectedly, are closely guarded by licensing companies.
Sharing a common aim of developing cost effective and efficient H2 PSA technology, HPCL
collaborated with Korean partner with an objective to set-up a commercial scale unit at one
of HPCL’s refineries. The project was undertaken synergizing the knowledge of HPCL and
its collaborator in different areas of engineering, materials and process know-how.
Subsequently, the first commercial H2 PSA unit for H2 purification from the Continuous
Catalytic Reformer (CCR) off-gas was installed successfully in HPCL Visakh Refinery.
PSA Process
Adsorption process is based on the phenomenon of attraction that a molecule from a gas
phase or liquid phase experience on the surface of a solid, named adsorbent. Adsorbent
materials are typically porous solids, having large specific surface area. The amount of
chemical species getting adsorbed onto the solid surface shall depend on several factors
such as fluid phase composition, temperature, pressure and most importantly, the
thermodynamic limit of the adsorbent material. The fundamental principle of PSA process is
adsorption of chemical species on the adsorbent surface at high pressure and release of the
same from the surface at lower pressure, hence the process is termed pressure swing
adsorption (PSA), the total pressure of the system “swings” between high pressure during
feed adsorption and low pressure during regeneration.
In case of purification of H2 using PSA, several layers of adsorbent materials are utilized for
selectively separating different impurities. Activated carbon can be used to remove impurities
such as hydrocarbons, moisture and CO2 quite selectively while Zeolite materials can be
utilized for removal of tougher / less adsorbed impurities such as N2 and CO. Also it has to
be kept in mind while selecting an adsorbent for PSA, that higher capacity always should not
be correlated with better separation process. It is often found that materials showing linear or
slightly nonlinear isotherms are preferred in PSA design, owing to the fact that regenerability
of such materials are superior.
From an engineering point of view, the most challenging part in this technology is the
development of cyclic strategies for effective pressure swinging to improve performance of
the PSA. Despite the performance of the material, the design of a PSA process requires
such engineering decisions which dictates performance of unit. It can be observed that the
step changes in a normal PSA are accomplished by the simultaneous operation of a
complex valve arrays connected to each adsorber vessels. Utilizing high performance and
precision valves and control philosophy, the opening or closing of valves are configured in a
PLC unit for controlling / operating the unit.
HPCL H2-PSA Plant at Visakh Refinery
The first commercial scale H2 PSA unit, developed by HPCL R&D Centre along with its
Korean collaborator, was successfully commissioned in its Visakh Refinery during May 2015.
The unit was designed for processing 6000 Kg/hr (36000 Nm3/hr) of CCR off-gas and
producing high purity H2 of 99.5% with high recovery.
The unit was designed with 6 adsorbers in service during normal operation. However, to
allow flexibility in any case of malfunctions, the unit has been designed for operation with
reduced number of adsorbers in service. The unit includes a compressor for increasing the
tail gas pressure from 0.5 kg/cm2g to about 4 kg/cm2g for transferring lighter hydrocarbons
into Refinery fuel gas network.
This effort of technology commercialization gave insights to HPCL about adsorbent selection
& loading, sequencing valve operation, and PLC programming which are critical to have a
sustained operation.
Adsorbent Replacement in PSA Plant of Mumbai Refinery
As a step to reinforce its capabilities in PSA technology, in a first-of-its kind effort, a project
was taken up by HPCL R&D for selection and replacement of adsorbents in one of its
existing commercial PSA unit in Mumbai Refinery. The PSA unit was designed to process
33400 Nm3/hr of feed gas coming from the steam reformer of HGU and to produce 99.99
mole% Hydrogen. The PSA was in service since its commissioning in 2000 and the
adsorbents have not been replaced since then. However, the Hydrogen recovery of the unit
has progressively deteriorated due to ageing of adsorbents.
HPCL R&D having acquired expertise in PSA technology including adsorbents, has carried
out research studies on numerous variety of adsorbents and selected the optimum
adsorbents. These were loaded in the PSA unit during the T&I. The unit was restarted
successfully and is presently operating meeting the design product purity and recovery.
The project involved the following stages.
a. Adsorbent Testing and Selection
It is most important factor which decides the performance of PSA unit. Testing of adsorbents
and their selection were done based on the following criteria.

Adsorption capacity: Generally termed as loading, this is one of the most important
characteristic of an adsorbent. It is the amount of adsorbate taken upon by the
adsorbent, per unit mass of the adsorbent. Also, capacity directly effects the capital
cost since it dictates the amount of adsorbent required, which also fixes the volume
of adsorber vessels. Hence, in this work, main focus was to identify the best suited
adsorbents in order to get superior product purity and increased hydrogen recovery.

Selectivity is the ratio of adsorption capacity of the adsorbent for one component to
that of another in a given fluid concentration. An ideal adsorbent would have very
high selectivity which means it adsorbs only the impurity without any uptake of the
major component.

Regenerability: All cyclic adsorption processes rely upon the regenerability, so that
the material can operate in sequential cycles while having the uniform performance.
The regenerability of an adsorbent affects the fraction of the original capacity that is
retained (working capacity) and time required for regeneration.

Cost of the adsorbent material.
To evaluate these parameters, available information from the existing databases, along with
vendor datasheets of the commercially available materials were used as a preliminary
shortlisting criteria. Further, by conducting in-house adsorbent testing and using PSA
simulation models, adsorbent materials were finalized.
b. Adsorbent Unloading
The spent adsorbent materials from the PSA vessels were removed after the plant was shut
down. Utilizing the vacuum unloading system and cyclone separators, the adsorbents were
safely removed and packaged for further safe disposal.
c. Vessel Cleaning and Inspection
Once the spent adsorbents were removed from the PSA vessels, the internals were cleaned
from any debris, corrosion or other impurities. Proper safety measures were taken for
working in the confined space for internal cleaning of vessels with wire brushes.
d. Adsorbent Loading
The type of adsorbent loading in a reactor or vessel system are determined by the type of
application and physical properties of the material to be utilized. Dense loading methodology
is generally favored since it enhances the quantity of material loaded in the system by about
10 to 15 % by reducing the bed void fraction. This will subsequently result in various benefits
in terms of higher throughput / more adsorbent life / better conversion / better flow
distribution. Alternately sock loading is performed when the adsorbent is physically weaker
and have tendency to form fines. The downside of sock loading being lower adsorbent
quantity/ higher probability of channelling/ maldistribution of flow causing non-uniform bed
profile for conversion or temperature/ higher attrition, and ultimately lower throughput.
One of the major challenges during the project execution was the loading methodology to be
adopted. By utilizing a ‘Static Dense Loader’ mechanism, HPCL was able to overcome the
loose packing of adsorbents in sock loading. A static mechanical device was used which
guided individual adsorbent particles with uniform spread while particles are allowed to freely
fall along the length of the column. The uniform dispersion ensured proper distribution of the
particles in the bed and helped in achieving much better density and homogeneity compared
with sock loading while overcoming the issues such as attrition and abrasion in comparison
with dense loading.
e. Controlled Loading
The type adsorbents selected for the HGU PSA plant presented with mainly two challenges.
Adsorbent loading in the humid atmosphere of Mumbai demanded sufficient blanketing for
the isolation of high capacity adsorbent materials from atmospheric moisture and other
impurities. Proper dust filtration system was also a pre-requisite for attending to fine dust
materials generated during the adsorbent loading activities. These difficulties were overcome
by utilizing N2 for creating inert atmosphere while adsorbents were loaded and a dedicated
high suction vacuum blower was used for controlling dust.
f. Unit Start-up
PSA unit was made ready for startup by carrying out high pressure leak tests and valve
sequencing tests. N2 blow out of vessels were done for removing any entrapped fine
particles from vessels. Once the HGU steam reformer was stabilised, PSA unit was taken
on-line and product H2 was routed to downstream units. Continuous monitoring of process
parameters were done until the PSA unit was stabilized.
g. Safe Practices
During the entire project execution safety was considered with paramount importance.
Matching to the industry standards, safe work practices, deployment of personal protective
equipment prior to entering the plant area, etc. was ensured through a dedicated safety
officer for entire duration of the project.
Conclusions
India has about 22 operating refineries with a capacity over 220 MMTPA. Hydrogen is
generated in refineries through steam reforming of methane or catalytic refining of naphtha.
PSA units are used for increasing the Hydrogen purity upto 99.99 mole% from a level of 70
mole% in order to utilize in Hydroprocessing units. Typically, each refinery has at least 2 or 3
PSA units for this purpose. In addition, growing hydrogen requirement is compelling refiners
to have additional PSA units for recovering Hydrogen from various refinery off gases
streams. This growing importance for Hydrogen in refineries has given lot of scope and
potential for establishing new H2 PSA units across the country.
Increasing hydrogen recovery in the existing PSA through adsorbent replacement adds
another credit to HPCL’s efforts for commercialization of indigenously developed H2 PSA
technology. By way of setting-up the commercial unit in Visakh Refinery, and successfully
revamping PSA unit of Mumbai Refinery, HPCL has now acquired significant experience, indepth knowledge and know-how on adsorbents, valves sequencing and PLC program, thus
making it a prospective supplier of H2 PSA technology.
_________________________________________________________________________
A statistical approach to tap hidden fuel saving potential in GTGs
Subhash Nandanwar, M.V.Borkar and Sardar Shaik
Hindustan Petroleum Corporation Limited - Visakh Refinery
1. INTRODUCTION:
Refining process is highly energy intensive. Energy cost of refining industry is as high as 5070% of total operating cost depending upon the configuration and the technology employed.
Power is one of the major and critical sources of energy required to operate refineries.
Reliability of power is a major concern across industries as it has substantial impact on the
bottom line of the business due to operating cost. Refining being integrated and continuous
process, it is always preferred to have Captive Power generation to ensure reliability.
Electrical energy constitutes about 30 - 40% of the total energy cost as GTs consume large
quantities of Naphtha or LNG. Thus, any small efforts towards reduction of Specific Fuel
Consumption (SFC) in GTGs will reduce the energy costs and thus add to the refining
margins.
Visakh Refinery has one Frame-3 (normally not in service) and four Frame-5 machines for
generating power with total design capacity of 94.8 MW. During normal run, Visakh
Refinery’s power requirement is in the range of 72-78 MW. Out of this, approximately 72-76
MW is generated from GTGs by firing naphtha and balance is taken from state grid.
2. SPECIFIC FUEL CONSUMPTION:
Specific fuel consumption of GTGs is the quantity of fuel consumed for generating unit
quantity of power (1 MWH). Basis the design heat rate of GTGs at Visakh Refinery, design
Specific Fuel Consumption (SFC) for Frame-3 GTG is approx. 0.382 MT/MWH at a load of
7.2 MW and 0.331 MT/MWH at a load of 21.9 MW for Frame-5 GTGs.
Over a period of 5 years, overall SFC of GTGs has improved from 0.363 MT/MWH to 0.345
MT/MWH due to the following:
Page 1 of 4


Optimization of GTGs operation i.e., shutting down Frame-3 GTGs and operating
Frame-5 GTGs on continuous basis.
Installation of magnetic resonators on fuel (naphtha) lines of all Frame-5 machines
which resulted in approximately 1% fuel savings.
Earlier all GTGs were operated in island mode and the load of individual GTG was
dependent on the load of process units connected to particular GTG. Hence, improvement in
overall SFC was limited. However, after implementation of load shedding scheme and
maximization of machine loads, further improvement was observed in SFC from 0.345 to
0.337 MT/MWH.
3. DATA ANALYSIS & INTERPRETATION:
To explore further possibility of improvement in SFC, past data of individual GTGs for a
period of 3 years was studied. This involved tabulating of load (MW) Vs actual SFC
(MT/MWH) for individual GTGs and developing the performance curves of each Frame-5
machine. These performance curves are obtained by plotting load, MW Vs actual SFC,
MT/MWH. Data points considered are more than 150 for each machine. Subsequently,
characteristic equations are developed by regression method, for estimating the specific fuel
consumption at various loads. These curves when plotted in a single graph are helpful in
comparing the performance of individual GTGs.
From the curves in the figure, following is inferred:


Specific fuel consumption of GTG-3 is lowest for all the load conditions followed by
GTG-4 for loads greater than 14.5 MW.
SFC for GTG-6 is lower than GTG-4 up to 14.5 MW load, thereafter it is more than
GTG-4.
Page 2 of 4

Similarly, SFC for GTG-5 is lower than GTG-4 up to 15.5 MW load, thereafter it is
more than GTG-4.
As can be seen from the above plot, the curves of individual GTGs indicate that though the
design SFC for all GTGs is same, it is different for each machine at varying load in actual
case. Thus, it is understood that there is further scope for improvement in SFC at the total
operating load requirement by optimizing loads on individual machines.
4. OPTIMIZATION OF LOADS:
To tap this hidden fuel saving potential and achieve improvement is SFC, it is required to
identify the optimum loads across individual GTGs. A simple excel solver function has been
used which accounts the characteristic equations of individual machine for arriving at the
lowest possible SFC for a given combined load requirement. Two cases have been worked
out using this solver function.
Case 1: All Frame-5 machines are running with a maximum load of 20 MW each as there are
no operating constraints on any of these machines at this load. The table -1 provides the least
possible SFCs which can be achieved for various load combinations basis excel solver
function.
Table-1
Total Load
requirement, MW
60
62
64
66
68
70
72
74
76
78
GTG-3
18
20
20
20
20
20
20
20
20
20
Load distribution, MW
GTG-4 GTG-5 GTG-6
20
11
11
20
11
11
20
11
13
20
11
15
20
11
17
20
11
19
20
12
20
20
14
20
20
16
20
20
18
20
SFC,
MT/MWH
0.334
0.330
0.329
0.327
0.325
0.323
0.322
0.321
0.320
0.318
From the above table, it can be seen that the least possible SFC even at a low load of 60 MW
is 0.334 MT/MWH.
From operating curve, the actual average SFC at a load of 20 MW for GTG-3, 4, 5 and 6
machines is 0.30, 0.310, 0.327 and 0.329 MT/MWH respectively. This indicates that SFC of
GTG-5 and 6 is higher as compared to GTG-3 and 4.
Case 2: Any 3 Frame-5 machines are running (assuming one machine is under CI) with the
same load constraints as considered in case 1. This case has been worked out to get the least
possible SFC from any three machines to meet the power requirement between 55 MW and
60 MW. The machine wise load distribution is provided in the table -2 below.
Table-2
Load distribution, MW
Total Load
SFC,
requirement, MW GTG-3 GTG-4
GTG-5
GTG-6 MT/MWH
55
20
20
15
S/D
0.317
55
20
S/D
15
20
0.324
55
S/D
20
15
20
0.327
Page 3 of 4
Total Load
requirement, MW
55
57
57
57
57
59
59
59
59
Load distribution, MW
GTG-3 GTG-4
GTG-5
GTG-6
20
20
S/D
15
20
20
17
S/D
20
S/D
17
20
S/D
20
17
20
20
20
S/D
17
20
20
19
S/D
20
S/D
20
19
S/D
20
20
19
20
20
S/D
19
SFC,
MT/MWH
0.318
0.315
0.322
0.326
0.316
0.314
0.320
0.324
0.314
The specific fuel consumption shown in the above tables is based on the naphtha flow meters
of individual machines. In order to verify the correctness of fuel flow meters, reconciliation
with naphtha tank gauging was done. The difference in fuel quantities between tank
measurement & flow meters is 1.1% which is within acceptable limits.
Basis above analysis and observations, following recommendations were provided and
implemented:
1. During four Frame–5 GTG operating scenario, maximize the load up to 76-78 MW
which will have the lowest possible SFC.
2. During low power requirement like in the range of 55 to 60 MW, it is advisable to run
GTG-III/IV/V and shutdown GTG-VI to achieve lowest possible SFC.
3. In all other cases, it is recommended to operate the GTGs as per the load distribution
given in the tables 1&2 for achieving best possible specific fuel consumption with the
existing set up.
After implementing this optimization program, lowest fuel consumption per unit power
generation in GTG’s could be achieved. SFC improved to 0.335 MT/MWH equivalent to
overall reduction of SFC by approx. 0.002 MT/MWH i.e., naphtha saving of ~0.6% per MW.
With this approx. 100 MT per month naphtha could be saved with monetary benefit of
approx. R.30 Lakhs.
5. CONCLUSION:
A simple statistical approach was used to identify the SFC of each machine by plotting the
curve against load, MW. This approach has helped in identifying the optimum load
distribution across individual machines for achieving the lowest possible SFC when the load
requirement is varying. Subsequently, load optimization was done and fuel saving of
approximately 100 MT per month could be achieved.
Page 4 of 4
Introduction of Hot Separator in Naphtha Hydrotreater
By: Anirban Ray / Debasis Sarma (Reliance Refinery Division)
Abstract:
Petroleum Refinery operates NHT units for preparing feedstock for Platformer. Naphtha
Hydrotreater (NHT) removes contaminants such as Olefins, Sulfur, Nitrogen and metals for
maintaining required Platformer feed spec. It was observed that energy efficiency of NHT can be
significantly improved by developing suitable energy conservation scheme. Installing Hot
Separator is a way to improve energy efficiency in Hydrotreaters.
In conventional design of Naphtha Hydrotreater, entire reactor effluent is directly routed to cold
separator through a series of Heat Exchangers and Reactor Effluent Air Cooler (REAC). However,
detailed study revealed the feasibility of introducing a Hot Separator in Reactor Effluent Loop,
which has high potential for energy saving since NHTs operate at relatively higher pressure.
Detailed study was carried out in-house using Aspen Plus, HTRI and Tray Hydraulic Software for
various configurations. Most optimum configuration was selected based on techno economic
evaluation.
Hot separator installed in NHT has provided energy benefit by reducing the rejection of heat in
Reactor Effluent Air Cooler. This heat saving in stripper bottom stream is used for steam
generation and to reduce fuel consumption in heater. Hot Separator scheme was successfully
commissioned in Naphtha Hydrotreater. Post commissioning of Hot Separator, significant
improvement in EII over previous year has been achieved. This is the first instance of installing
Hot Separator in NHT for the particular technology supplier that has been commissioned
successfully in a commercial unit. Unit is in continuous operation post implementation of Hot High
Pressure Separator (HHPS) scheme for substantial period and operating smoothly with sustained
benefit.
Introduction
Naphtha Hydrotreater in Petroleum Refineries removes contaminants such as Olefins, Sulfur,
Nitrogen and metals for maintaining required Platformer feed spec. NHT is a major energy
consumer in energy intensive refining business. In a typical NHT, major energy consumers are
Naphtha Splitter Reboiler (provides duty to separate full range naphtha into light and heavy
naphtha fractions), Charge Heater and Stripper Reboiler, whereas heat is rejected into
atmosphere by Reactor Effluent Air Cooler and Stripper overhead condenser Air Cooler.
Opportunity for Energy Consumption
In order to minimize overall energy, energy consumption in NSPL Reboiler, Stripper Reboiler, and
Charge Heater are to be minimized, and heat rejected from Reactor Effluent Air Cooler (REAC)
and Stripper overhead condenser to be reduced.
NHT Reactor effluent stream is cooled in CFE (Combined Feed Effluent Exchanger) and REAC
from very high temperature to almost ambient temperature and subsequently Separator liquid is
again reheated to Stripper feed temperature, utilizing heat available in Stripper bottom. If Stripper
feed temperature can be maintained without using heat from Stripper bottom and thereby
reducing heat rejected in REAC, energy available in Stripper bottom can be utilized in much better
energy efficient way. If HNUU separator pressure is high enough (for targeting high HDN), there
exists an opportunity for introducing a hot separator at CFE downstream, as significant amount
of hydrotreated liquid gets separated in Hot High Pressure Separator, which can then be directly
routed to Stripper, keeping Stripper bottom heat available for use.
Available heat from the Stripper Reboiler can be utilized in following areas:
1. Naphtha Splitter Reboiler duty (saving steam / fuel)
2. Feed preheating (saving fuel consumption in Charge Heater)
3. Medium Pressure Steam Generation at Stripper Bottom
4. Naphtha Splitter cold feed preheating
5. Supplying NHT bottom at high temperature to Platformer (energy saving in Platformer)
Looking at Refinery steam balance, MP Steam generation is always beneficial. As stripper bottom
temperature is high enough for MP Steam Generation, option 3 was adopted. While developing
the scheme, emphasis were given on the following points:
1. No modification in recycle gas compressor
2. No new exchanger in Recycle gas / high pressure loop for eliminating safety risks
3. Maximize utilization of existing exchanger
Description of the modification:
NHT Reactor effluent ex CFE is routed to new Hot Separator. Liquid separated in Hot Separator
is routed to Stripper. Hot Separator liquid has provisions to be routed through preheat exchangers,
where it is heated by Stripper bottom liquid. Hot Separator vapor, after mixing with Net gas (makeup Hydrogen) and wash water is routed to Cold Separator through REAC. Provision exists for
wash water injection at Hot Separator inlet. Cold Separator liquid after preheating by Stripper
bottom, joined with Hot Separator liquid and routed to Stripper at desired Stripper feed
temperature. Stripper bottom, after generating MP Steam and Hot Separator / Cold Separator
liquid preheating, is sent to Platformer. MP Steam generated is mostly used within the unit though
provision exists for export to steam network.
Salient features of the scheme:
Installation of Hot Separator is a major revamp in NHT with significant modifications in process
flow scheme. This modification also involves change in hydraulic load, equipment adequacy and
impact on process reliability. The major impact of installing Hot Separator in NHT are described
below in detail:
Finalizing Hot Separator temperature: Technically, higher the Hot Separator temperature, better
it is for energy efficiency, as higher Hot Separator temperature indicates higher Stripper feed
temperature, resulting in reduction in reboiler duty. However, if Stripper feed temperature
becomes too high, there is a possibility that most of the heat is lost through Stripper overhead,
additionally, adequacy of Stripper overhead condenser is also required to be checked. Hot
Separator temperature also determines extent of liquid separation and consequently hydrocarbon
flow to Cold Separator through REAC. Higher the Hot Separator temperature, lower the liquid
separation in Hot Separator and consequently higher flow to REAC and Cold Separator. Hence
to arrive at the most energy efficient point, a trade-off is required between Hot Separator
temperature and Hot Separator liquid generation. In addition, detail evaluation of CFE is required
to estimate feasibility of achieving desired Hot Separator temp. If wash water is required to be
injected upstream of Hot Separator, higher temperature will require increased wash water
addition, whereas lower Hot Separator temperature will require reduced wash water injection.
Wash water injection requirement is to be cross-checked against available wash water quantity.
Hydrogen Loss: Due to reverse-order solubility Hot Separator liquid is rich in Hydrogen and results
in higher solution loss. After installation of Hot Separator, Hydrogen solution loss is expected to
increase slightly. Due to reduced sponging effect in the Cold Separator, less C1, C2s will go into
Cold Separator liquid, resulting in reduced lighters (C1 / C2) in Stripper off gas. Consequently
Stripper off gas will become lighter.
Recycle gas purity: Due to less sponging effect in Cold Separator after installation of Hot
Separator, less lighter hydrocarbon (C1/C2) will be generated in Stripper overhead and hence
Recycle gas will have more lighter Hydrocarbon (C1/C2) with less Hydrogen (due to high solution
loss) and high molecular weight. Recycle gas molecular weight will increase slightly and
consequently Hydrogen purity in recycle gas will drop in same order.
Pressure drop in Reactor effluent circuit: Installation of Hot High Pressure Separator reduces a
significant amount of Hydraulic Load in Hot Separator Overhead Circuit, as majority of
hydrotreated liquid product is removed in the Hot Separator itself.
Equipment adequacy: New Hot Separator and associated exchangers has been designed based
on finalized Hot Separator temperature. Existing Stripper adequacy has been evaluated, as vapor
liquid load in Stripper will change post installation of Hot Separator. Duty of REAC (Reactor
Effluent Air Cooler) is expected to be reduced significantly post Hot Separator installation. In some
cases, few bays of air coolers may require to be isolated. To maintain velocity through air cooler
tubes within recommended range remains always a challenge post Hot Separator installation.
Process reliability: Hot Separator temperature also governs reliability of the unit as slippage of
ammonium chloride salt into Stripper depends on Hot Separator temperature. If Hot Separator
temperature is higher, probability of ammonium chloride slippage into Stripper decreases and if
Hot Separator temperature is lower, probability of ammonium chloride slippage into Stripper
increases. Sometimes, when chloride (in feed / make-up gas) is high, it is advised to inject wash
water at inlet of Hot Separator to eliminate possibility of ammonium chloride ingress into the
Stripper. However wash water addition impacts the flow scheme in several ways as following:
 Hot Separator design may change
 Hot Separator temperature will reduce – thus reducing the benefit on account of energy
saving.
 Stripper adequacy is to be checked as routing of some equilibrium water into Stripper will
affect Stripper vapor liquid load significantly.
New Equipment:
Following new major pieces of equipment were installed for the selected scheme, in addition to
utilizing existing resources:
1. Hot High Pressure Separator
2. New MP Steam generator
3. New Cold Separator / Hot Separator liquid preheater
Simplified Sketch
Recycle Gas
Compressor
Recycle
Gas
Hydrogen
Make-up
Feed
Feed Surge Drum
Pump
CFE
Charge
Heater
Reactor
REAC
Wash
Water
1
Stripper Off-Gas
Stripper
Hot
Separator
Cold
Separator
Hot
Separator
PreHeater
Cold
Separator
PreHeater
3
Reboiler
Furnace
MP Steam
To Platformer
MP
Steam
generator
2
BFW
1
Hot High Pressure Separator
2
MP Steam Generator
3
Cold Separator / Hot Separator liquid preheater
Benefit of Proposed Modification
Indigenously developed EnCON Scheme for installation of Hot High Pressure Separator in reactor effluent
circuit was successfully commissioned in NHTs. Installation of Hot Separator allowed margin to trim down
heat rejection in REAC (Reactor Effluent Air Cooler) and utilize recovered heat in stripper bottom circuit
to produce MP Steam.
Table: Estimated Benefit
UoM
Configuration without Hot
Separator
Configuration with Hot
Separator
Net Gain due to saving in
Stripper Reboiler energy
Gcal/hr
Base
Base + 6
MP Steam Generation
Gcal/hr
Base (Zero)
Base + 8.7
Nm3/MT
Base
Base + 1.3
m3/h
Base
Base
%
Base
Base -30%
Parameter
Loss in Hydrogen Solubility
through stripper off gas
Increase in Wash Water
Energy Intensity Index
Conclusion:
Installation of Hot Separator in NHT is an innovative scheme, with aim for achieving significantly
higher energy efficiency in NHT. From conceptualization to commissioning, this in-house project
has been successfully implemented after detailing by licensor while achieving desired benefit with
assured integrity and reliability. As highlighted in previous section, owing to this project, Energy
Intensity Index (EII) of NHT has been significantly improved. Smooth commissioning and
subsequent sustained operation is a testament to flawless design and execution of this project.
Case study for
Refinery Technology Meet
under the category
Poster Presentation
INNOVATIVE WAY TO REPLACE
INTERNAL GASKET OF CCR
PLATFORMING REACTOR REDUCES
UNIT DOWNTIME AND SET NEW
BENCHMARK
BHARAT OMAN REFINERIES LIMITED
Administrative Building, Refinery Complex, Post BORL Residential Complex, Bina – 470124,
District - Sagar, Madhya Pradesh, India.
Tel: +917580226000, Fax: +917580226903, Website: www.borl.in, Email: [email protected]
Application
1.
Name of the Technology/Process/Product
Innovative way to reduce the CCR unit downtime
2.
Name of individuals, orgonisation
Akhilesh Kumar saxena
Ashish kumar
Manish Pandey
Rakesh Sharma
Vikant Maithil
Deepesh Kumar
Bharat Oman Refineries Limited (BORL), Bina, Madhya Pradesh
ABSTRACT
There are various innovations being carried out in the industries to improve production
and productivity.
At BORL by using innovative method the CCR unit reactors gaskets were replaced with
33% of time against internationally available benchmarks. This resulted in avoiding
revenue loss of INR 30Cr on production downtime apart from heavy rigging resource
mobilization worth INR 2Cr.
This paper discusses conventional as well as alternate approach to replace the CCR
reactor center pipe support gaskets in Reactor-01
Conventional approach requires involvement of huge manpower, time, and efforts and
also necessitates disturbing reactor internals .Where as applying innovation gasket
replacement job was completed in much lesser time. The knowledge & experience
gained being replicated across industries.
1. INTRODUCTION
BORL is equipped with a CCR unit. The unit is technology is from a reputed licenser. The
unit takes treated Heavy Naphtha as feed stock and employs platinum based catalyst.
The objective of the unit is to maximize Aromatics by converting Naphthenes and
Paraffin into Aromatics. The reaction of Naphthenes and Paraffin’s into Aromatics
produces Hydrogen as a by-product, known as Hydrogen Rich gases.
Naphtha's (typically having low octane ratings) converted into high-octane liquid products
called reformates through Reactors. In these reactors catalyst flows by gravity out from
Reactor bottom to the top of the next reactor through catalyst Transfer pipe. Catalyst flow
between the reactors is through equally-spaced transfer lines designed to ensure uniform
catalyst flow through the catalyst bed. We are having 3 reactors 20-RB-00-101/102/103
in MSB at Bharat Oman refineries Limited and to ease catalyst transfer by gravity, CCR
catalyst handling vessels are stacked in nearly vertical line in the structure.
The problem of Center pipe support gasket rupture was encountered in Reactor 1
leading to plant shutdown. To minimize the downtime the job of Gasket replacement was
carriedout in an innovative fashion setting a new benchmark for industry.
2. OBSERVATIONS
On 24rd April, 2016 Unit faced the sudden major upsets as mentioned below which led to
unit shutdown





Reactor # 2 differential pressure increased rapidly to 0.7 Kg/cm² (g) – Reading
went out of Range.
The delta pressure (ΔP) between Reduction Zone & Rx#1 dropped to Zero in 510 minutes.
Reactor # 3 delta temperature (ΔT) increased.
RG Compressor tripped due to high back pressure.
Reactor # 1 & Reactor # 3 remained pretty much stable.
DCS snap shots for the unit parameter trend prior to unit shutdown
The trends of parameters were indicating towards the damage /disturbance of reactor
internal components.
It was analysed that following are likely to be probable reasons behind the process upset





Opening of any Duro lock coupling
Damage in Scallop pipe/center pipe
Damage in mitre elbow
Damage in Center pipe support
Damages in any of the sealing gaskets
Upon Reactor Opening, Initial inspection revealed Center Pipe Support (CPS) to
Rx#1bottom Dish End Gasket joint damaged / failed, leading to loss of Catalyst
Containment from Rx#1 Reduction Zone Drum and significant gap in Center Pipe
Support (CPS) to Rx#2 bottoms Dish End Gasket joint.
The failure of gasket was attributed to MCC (Metal catalysed coking) due to prolonged
operation.
Center Pipe Support (CPS) to Rx#1bottom Dish End Gasket joint damaged / failed
3. LIMITATION FACED AND ANALYSIS
The standard method (Annexture-1) of Center pipe support gasket replacement
required Removal of reduction zone and reactor internal weighing more than 15 tonne at
60 m height which has following limitations
1. Time required for conventional method was 25 days even when meticulously
planned
2. Emergency heavy equipment mobilization would have required additional 5day
BORL being a single stream refinery the cumulative outage of 30days including 26
engineering days
( Schedule Annex-3) would have led to completely exhausting
intermediate tanks leading to refinery shut down . This would have affected the
company’s product supply chain& brand apart from revenue loss.
Impending situation posed a challenge which led team to look into possibility of
completing job with available resources within shortest possible time.
The team formed the following strategy
1. Analysing the limitation of conventional method
2. Analyse each of the reactor component
3. Find Possible ways to overcome the limitation
4. Provide the alternate with substance
3.1 Followings were the findings
The two largest time & resources consuming activities were


Lifting of reduction zone (Wt. 10T)
Lifting reactor 1 internals (Wt. 2T)
This was required due to following design constraint
 Center pipe support gasket size was larger than man-hole size
 There was no provision to place the gasket in position without removing of center
pipe which requires removal of reduction zone
The team analysed the reactor components design and explore the possibility of
overcoming design limitations.
During the analysis of problem and reactor internal components it is found
1. Gasket is not a pressure bearing component
2. A bellow is provided at top of center pipe with axial compression limit of 40mm
3.2 The Innovative Way Forward
Considering above findings following key changes in conventional methods were
adopted in consultation with unit licensor
1. To cut the gasket in four pieces to make possible to put through manhole
2. To shift the load of center pipe from center support to reduction zone lifting lugs
through chain pulley blocks
3. Arrangements to be made to lift the center support internally to insert gasket
A detail procedure was developed considering the above three key points with schedule
of 8 engineering days ( Annexure -4) . The activites marked with yellow in Convention
schedule (Annex-3) were eliminated .
3.3 The Execution with innovation
Following were the innovation of the methods

Available Spare Gasket cut into four pieces through laser cutting

Load Shifting Arrangement
compression limit
from
center
support
within
bellow
Placement of lifting arrangement on the center pipe lugs. 4 chain pulleys of 2
tons each with equal run of long travel were placed to lift the center pipe by
using the lifting lugs of reduction zone bottom dish ends. The center pipe was
lifted by 12mm using bellow compression with in the Bellow Compression limit
of 40mm. This Job was performed with precision monitoring compression by
Dial gauge.
4 Chai pulley lo k of e ual u
pla ed fo lifti g the e te pipe
Schematic diagram of center pipe lifting arrangements
Lifting of center support for inserting cut gasket
After shifting the load of center pipe on chain pulleys . The center pipe support plate
was lifted uniformly using long Hold down stud to insert the gasket
Schematic diagram of center pipe support
The assembly was done in reverse order
4.






CONCLUSION
The job was executed within 8 engineering days against estimated time of 26 days
The 19 major activities were eliminated in innovative schedule
The cumulative downtime was reduced to 14days from estimated 30~32 days
The revenue loss of 20 days’ worth INR 30 Crore was averted
The plant is running normal for last 4 months
The same innovative method was replicated at another PSU refinery Successfully
when it faced similar failure
Annexure -1
Conventional procedure for center pipe support gasket replacement in CCR
reactor
Rx2 gasket replacement requires
 Remove reduction zone and associated accessories
 Remove all internal of Rx-01 and Rx02
Rx1 gasket replacement requires
 Remove reduction zone and associated accessories (Common for both case)
 Remove all internal of Rx-01
Reactor internals which are to be removed before perform the gasket replacement
job
First step to dismantle all the dur-o-lock couplings inside the reactor after that remove all
the internals as per following:
 Cover deck plates (Wedge pins/anchor chains )
 Miter elbow
 Center pipe with expansion bellow
 Support bottom plate
Step wise procedure to replacement of center pipe support gasket
Step1. Remove the reduction zone (Approx 10ton) and associated accessories with
the help of crane
Step2. Dismantle all the dur-o-lock coupling (9 no’s in Rx#01 and 12 no’s in Rx#2)
Step3. Remove the miter elbow of Rx#01 and cover deck plates
Step4.Now remove reactor center pipe along with expansion bellow with the help of
crane and put on ground.
Step5. Remove all the bolts associated with center pipe support (Hot work may be
required to cut the bolts M24XX180 B16 and nut Gr.3) thereafter, loose the hold down
bolt and make arrange to lift the bottom support.
Step6. At last remove bottom support plate.
Now remove all the internals of Rx#02 in same sequence through Rx#01 and
reduction zone
Step1. Dur-o-lock coupling dismantle
Step2. Reduction zone removing
Height - Apr.
6.8M
Diameter - 920mm
Step3. Miter elbow removing from its
Step4. Expansion bellow
Step4. Remove center pipe along with expansion
Step 5&6. Removal of center pipe support bottom
Step wise activites to replace the center pipe gasket through conventional procedure
Step7. After removing all the internals, clean the gasket seating face in CCR reactor #2 ,
put new gasket and box up all the internals in reverse sequence .
Step8. After completion of seventh step replace the gasket same way in Rx-1 and install
all the internals.
Step9. Reduction zone box up along with new gasket and put all the necessary spool
pieces nucleonic gauges, electric heaters etc., on their position.
Annexure -2
INNOVATIVE METHOD
Step wise procedure to replacement of centre pipe support gasket
Step1 Dismantle all the dur-o-lock coupling (9 no’s in Rx#01 and 12 no’s in Rx#2)
Step2. All Deck plate and catalyst transfer pipes were removed opened for providing
access for movement and placement of lifting arrangement on the center pipe lugs. 4
chain pulleys of 2 tons each with equal run of long travel were placed to lift the center
pipe by using the lifting lugs of reduction zone bottom dish ends.
4 Chai pulley lo k of e ual u
pla ed fo lifti g the e te pipe
Schematic diagram of center pipe lifting arrangements
Step3. Opened all the head bolts of the center pipe support plate for gasket
replacement (44 bolts) from top of Reactor 2 and Cutting the seized bolts with nuts in
the top of reactor 2 using grinding machine. Cutting of head of the still seized bolts
from the bottom of reactor 1 and Using the specially modified Drilling machine to drill
out the seized bolts.
Schematic diagram of center pipe support
Step4. Lift of the center pipe with its support restricting the lifting gap not more then
12 MM (Minimum gap required for replacement of the gaskets) rest the center pipe
support on wedges by lowering the chain pulleys after removal of the damaged gasket
so to remove the load from the stud that is connecting the center pipe support and
center pipe.
Caution: As there is no free movement parts so center pipe displacement shall not
exceed maximum axial compression (20mm) of expansion bellow. Linear and
lateral displacement kept on observation by the help of dial gauge and spirit level
and it should be avoided.
Internal view of expansion bellow installed inside the reactor
Step5. Now remove old gasket and put new gasket in reactor 1 through top of reactor
2
A gap of
as ade y
lifti g the e te pipe fo
pla i g the gasket
Gasket i fou pie es
ith do e-tail joi ts as
pla ed f o top of
ea to
Schematic diagram of access to support bottom gasket
Step5. Placement of cut gasket having Dock tail joints with precision from Top of
reactor 2 and put the graphite tape adhesive on all the joints for filling any gap in
joints. Provide the 6 mm Ceramic rope between the center pipe support guide and
center pipe to assuring the complete sealing and avoiding the metal to metal contact.(
The gap should be limited in between 1-4 mm only)
Step6. Release the center pipe from the chain pulleys Checking and correcting the
verticality by using plumb of the center pipe(allowed tolerance ±19 MM) and Box up
of transfer pipe with deck plates
Schematic diagram of center pipe support gasket
Repeat the same activity for gasket replacement of the other reactors.
Annex-3
Attached PDF file
Annex-4
Attached PDF file
Conventional Approch For Central Support For Gasket Replacement
#
Activity ID
Activity Name
1
Original Start
CCR Reactor
R
Conventional Approch For Cen
Finish
Classic Schedule Layout with Line Number
Apr 17
Apr 24
May 01
M T W T Fri S S M T W T Fri S
Duration
26.0d 23-Apr-16 12:00 AM
19-May-16 12:00 AM
May 08
30-Aug-16 12:25 PM
May 29
S M T W T Fri S S M T W T Fri S S M T W
May 15
S M T W T Fri S S M T W T Fri S
May 22
19-May-16 12:00 AM, CCR Reactor Conventional
2
A1000
Blinding of Reactor
0.5d 23-Apr-16 12:00 AM
23-Apr-16 12:00 PM
Blinding of Reactor
3
A1010
Catylyst unloading
1.5d 23-Apr-16 12:00 PM
25-Apr-16 12:00 AM
4
A1020
Manhole Openning (for problem identification)
0.5d 25-Apr-16 12:00 AM
25-Apr-16 12:00 PM
5
A1040
Removal of control valves,electrical heaters from top of rea
1.0d 25-Apr-16 12:00 AM
26-Apr-16 12:00 AM
6
A1080
Top structure Cutting for reduction zone removal
0.5d 25-Apr-16 12:00 AM
25-Apr-16 12:00 PM
7
A1090
Reduction zone girth flange openning
1.5d 25-Apr-16 12:00 AM
26-Apr-16 12:00 PM
8
A1100
Placement of 300 ton crane
0.5d 26-Apr-16 12:00 PM
27-Apr-16 12:00 AM
9
A1110
Opening of Catalyst transfer pipes
0.5d 25-Apr-16 12:00 PM
26-Apr-16 12:00 AM
10
A1120
Removal of reduction zone and putting it on ground floor
0.5d 26-Apr-16 12:00 PM
27-Apr-16 12:00 AM
11
A1130
Opening and removal of mitter Elbow of RX1
0.5d 27-Apr-16 12:00 AM
27-Apr-16 12:00 PM
12
A1140
Removal bellow with frame after locking of RX1
0.5d 27-Apr-16 12:00 PM
28-Apr-16 12:00 AM
13
A1150
Entry enside central pipe for opening of Hex nut of Rx1
0.5d 28-Apr-16 12:00 AM
28-Apr-16 12:00 PM
14
A1160
Removal of center pipe of Rx1
0.5d 28-Apr-16 12:00 PM
29-Apr-16 12:00 AM
15
A1170
Opening of Center pipe support bolts of Rx1
0.5d 29-Apr-16 12:00 AM
29-Apr-16 12:00 PM
Opening of Center pipe support bolts of Rx1
16
A1180
provision of cutting of bolts(if not opened) of Rx1
0.5d 29-Apr-16 12:00 AM
29-Apr-16 12:00 PM
provision of cutting of bolts(if not opened) of Rx1
17
A1190
removal of center pipe support of Rx1
0.5d 29-Apr-16 12:00 PM
30-Apr-16 12:00 AM
18
A1200
opening and removal of mitter Elbow of RX2
0.5d 30-Apr-16 12:00 AM
30-Apr-16 12:00 PM
19
A1210
Removal bellow with frame after locking of RX2
0.5d 30-Apr-16 12:00 PM
01-May-16 12:00 AM
20
A1220
Entry enside central pipe for opening of Hex nut of RX2
0.5d 01-May-16 12:00 AM
01-May-16 12:00 PM
21
A1230
Removal of center pipe of RX2
0.5d 01-May-16 12:00 PM
02-May-16 12:00 AM
22
A1240
Opening of Center pipe support bolts of RX2
0.5d 02-May-16 12:00 AM
02-May-16 12:00 PM
Opening of Center pipe support bolts of RX2
23
A1250
provision of cutting of bolts(if not opened) of RX2
0.5d 02-May-16 12:00 AM
02-May-16 12:00 PM
provision of cutting of bolts(if not opened) of RX2
24
A1260
removal of center pipe support of RX2
0.5d 02-May-16 12:00 PM
03-May-16 12:00 AM
25
A1270
Cleaning of Flange face of RX2
0.5d 03-May-16 12:00 AM
03-May-16 12:00 PM
26
A1280
placement of center support gasket(oval) of RX2
0.5d 03-May-16 12:00 PM
04-May-16 12:00 AM
27
A1290
Putting back center support by matching the gaskets hole o
0.5d 04-May-16 12:00 AM
04-May-16 12:00 PM
28
A1300
TorqueTightening of center support pipe flange at 4.2 Nm
1.0d 04-May-16 12:00 PM
05-May-16 12:00 PM
29
A1310
Placing of Center pipe on support plate of Rx2
0.5d 05-May-16 12:00 PM
06-May-16 12:00 AM
30
A1320
tightenning of Center pipe hexbolt and filling of ceramic as
0.5d 06-May-16 12:00 AM
06-May-16 12:00 PM
31
A1330
catalyst Transfer pipe(duro lock) coupling boxup of Rx2
0.5d 06-May-16 12:00 PM
07-May-16 12:00 AM
catalyst Transfer pipe(duro lock) coupling boxup of Rx2
32
A1340
Alignment of center pipe (limits +- 19 mm) of RX2
0.5d 07-May-16 12:00 AM
07-May-16 12:00 PM
Alignment of center pipe (limits +- 19 mm) of RX2
33
A1350
placing of Bellow of RX2
0.5d 07-May-16 12:00 PM
08-May-16 12:00 AM
34
A1360
torquing of Elbow flange of RX2
1.0d 08-May-16 12:00 AM
09-May-16 12:00 AM
35
A1370
Cleaning of Flange face of RX1
0.5d 09-May-16 12:00 AM
09-May-16 12:00 PM
36
A1380
placement of center support gasket(oval) of RX1
0.5d 09-May-16 12:00 PM
10-May-16 12:00 AM
37
A1390
Putting back center support by matching the gaskets hole o
0.5d 10-May-16 12:00 AM
10-May-16 12:00 PM
38
A1400
TorqueTightening of center support pipe flange at 4.2 Nm
1.0d 10-May-16 12:00 PM
11-May-16 12:00 PM
39
A1410
Placing of Center pipe on support plate
0.5d 11-May-16 12:00 PM
12-May-16 12:00 AM
40
A1420
tightenning of Center pipe hexbolt and filling of ceramic as
0.5d 12-May-16 12:00 AM
12-May-16 12:00 PM
41
A1430
Alignment of center pipe (limits +- 19 mm) of RX1
0.5d 12-May-16 12:00 PM
13-May-16 12:00 AM
42
A1440
placing of Bellow of RX1
0.5d 13-May-16 12:00 AM
13-May-16 12:00 PM
43
A1450
putting back mitter elbow of RX1
0.5d 13-May-16 12:00 PM
14-May-16 12:00 AM
44
A1460
torquing of Elbow flange of RX1
0.5d 14-May-16 12:00 AM
14-May-16 12:00 PM
45
A1470
torque tightenning of reduction zone girth flange
1.5d 14-May-16 12:00 PM
16-May-16 12:00 AM
46
A1480
catalyst Transfer pipe(duro lock) coupling boxup
0.5d 16-May-16 12:00 AM
16-May-16 12:00 PM
47
A1490
manway box up
0.5d 16-May-16 12:00 PM
17-May-16 12:00 AM
48
A1500
Catalyst loading and Deblinding
2.0d 17-May-16 12:00 AM
19-May-16 12:00 AM
Catylyst unloading
Manhole Openning (for problem identification)
Removal of control valves,electrical heaters from top of reactor
Top structure Cutting for reduction zone removal
Reduction zone girth flange openning
Placement of 300 ton crane
Opening of Catalyst transfer pipes
Removal of reduction zone and putting it on ground floor
Opening and removal of mitter Elbow of RX1
Removal bellow with frame after locking of RX1
Entry enside central pipe for opening of Hex nut of Rx1
Removal of center pipe of Rx1
removal of center pipe support of Rx1
opening and removal of mitter Elbow of RX2
Removal bellow with frame after locking of RX2
Entry enside central pipe for opening of Hex nut of RX2
Removal of center pipe of RX2
removal of center pipe support of RX2
Cleaning of Flange face of RX2
placement of center support gasket(oval) of RX2
Putting back center support by matching the gaskets hole of RX2
TorqueTightening of center support pipe flange at 4.2 Nm of RX2
Placing of Center pipe on support plate of Rx2
tightenning of Center pipe hexbolt and filling of ceramic asbestos 6 mm rope
placing of Bellow of RX2
torquing of Elbow flange of RX2
Cleaning of Flange face of RX1
placement of center support gasket(oval) of RX1
Putting back center support by matching the gaskets hole of RX1
TorqueTightening of center support pipe flange at 4.2 Nm of RX1
Placing of Center pipe on support plate
tightenning of Center pipe hexbolt and filling of ceramic asbestos 6 mm rope
Alignment of center pipe (limits +- 19 mm) of RX1
placing of Bellow of RX1
putting back mitter elbow of RX1
torquing of Elbow flange of RX1
torque tightenning of reduction zone girth flange
catalyst Transfer pipe(duro lock) coupling boxup
manway box up
Catalyst loading and Deblinding
Note: Activites highlighted as yellow are eliminated or nor performed in innovative approch
Actual Level of Effort
Remaining Work
Milestone
Actual Work
Critical Remaining Work
summary
Page 1 of 1
TASK filter: All Activities
© Oracle Corporation
innovative Approch
#
Activity ID
1
Activity Name
Original Start
CCR REACTOR2
R
innovative Approch
Finish
Classic Schedule Layout with Line Number
Apr 17
Apr 24
May 01
M T W T Fri S S M T W T Fri S
Duration
10.4d 23-Apr-16 12:00 AM
03-May-16 09:36 AM
May 08
S M T W T Fri S S M T W T Fri S
31-Aug-16 07:07 PM
May 29
S M T W T Fri S S M T W T Fri S S M T W
May 15
May 22
03-May-16 09:36 AM, CCR REACTOR2 innovative Approch
2
A1000
Blinding of Reactor
0.5d 23-Apr-16 12:00 AM
23-Apr-16 12:00 PM
Blinding of Reactor
3
A1010
Catylyst unloading
1.0d 23-Apr-16 12:00 PM
24-Apr-16 12:00 PM
4
A1020
Manhole Openning (for problem identification)
0.3d 24-Apr-16 12:00 PM
24-Apr-16 07:12 PM
5
A1030
Opening of Catalyst transfer pipes of Rx1
0.3d 24-Apr-16 07:12 PM
25-Apr-16 02:24 AM
6
A1040
Placement of chain pulley blocks for lifting center support p
0.3d 25-Apr-16 02:24 AM
25-Apr-16 09:36 AM
7
A1060
Opening of Center pipe support bolts of Rx1
0.5d 25-Apr-16 09:36 AM
25-Apr-16 09:36 PM
8
A1070
provision of cutting of bolts(if not opened) of Rx1
1.0d 25-Apr-16 09:36 PM
26-Apr-16 09:36 PM
9
A1080
Lifting of center pipe support of Rx1 by chain blocks
0.2d 26-Apr-16 09:36 PM
27-Apr-16 02:24 AM
Lifting of center pipe support of Rx1 by chain blocks
10
A1100
Cleaning of Flange face of RX1
0.2d 27-Apr-16 02:24 AM
27-Apr-16 07:12 AM
Cleaning of Flange face of RX1
11
A1110
placement of center support gasket(oval) of RX1
0.2d 27-Apr-16 07:12 AM
27-Apr-16 12:00 PM
placement of center support gasket(oval) of RX1
12
A1120
Putting back center support by matching the gaskets hole o
0.1d 27-Apr-16 12:00 PM
27-Apr-16 02:24 PM
Putting back center support by matching the gaskets hole of RX1
13
A1130
TorqueTightening of center support pipe flange at 4.2 Nm
0.5d 27-Apr-16 02:24 PM
28-Apr-16 02:24 AM
14
A1140
Placing of Center pipe on support plate
0.2d 28-Apr-16 02:24 AM
28-Apr-16 07:12 AM
Placing of Center pipe on support plate
15
A1150
tightenning of Center pipe hexbolt and filling of ceramic as
0.1d 28-Apr-16 07:12 AM
28-Apr-16 09:36 AM
tightenning of Center pipe hexbolt and filling of ceramic asbestos 6 mm rope
16
A1160
Alignment of center pipe (limits +- 19 mm) of RX1
0.1d 28-Apr-16 09:36 AM
28-Apr-16 12:00 PM
Alignment of center pipe (limits +- 19 mm) of RX1
17
A1170
Placement of chain pulley blocks for lifting center support p
0.3d 28-Apr-16 12:00 PM
28-Apr-16 07:12 PM
18
A1180
Opening of Center pipe support bolts of RX2
0.5d 28-Apr-16 07:12 PM
29-Apr-16 07:12 AM
19
A1190
provision of cutting of bolts(if not opened) of RX2
1.0d 29-Apr-16 07:12 AM
30-Apr-16 07:12 AM
20
A1200
Lifting of center pipe support of Rx2 by chain blocks
0.2d 30-Apr-16 07:12 AM
30-Apr-16 12:00 PM
Lifting of center pipe support of Rx2 by chain blocks
21
A1210
Cleaning of Flange face of RX2
0.1d 30-Apr-16 12:00 PM
30-Apr-16 02:24 PM
Cleaning of Flange face of RX2
22
A1220
placement of center support gasket(oval) of RX2
0.2d 30-Apr-16 02:24 PM
30-Apr-16 07:12 PM
placement of center support gasket(oval) of RX2
23
A1230
Putting back center support by matching the gaskets hole o
0.1d 30-Apr-16 07:12 PM
30-Apr-16 09:36 PM
Putting back center support by matching the gaskets hole of RX2
24
A1240
TorqueTightening of center support pipe flange at 4.2 Nm
0.5d 30-Apr-16 09:36 PM
01-May-16 09:36 AM
TorqueTightening of center support pipe flange at 4.2 Nm of RX2
25
A1250
tightenning of Center pipe hexbolt and filling of ceramic as
0.1d 01-May-16 09:36 AM
01-May-16 12:00 PM
tightenning of Center pipe hexbolt and filling of ceramic asbestos 6 mm rope
26
A1260
Alignment of center pipe (limits +- 19 mm) of RX2
0.1d 01-May-16 12:00 PM
01-May-16 02:24 PM
Alignment of center pipe (limits +- 19 mm) of RX2
27
A1270
catalyst Transfer pipe(duro lock) coupling boxup
0.3d 01-May-16 02:24 PM
01-May-16 09:36 PM
catalyst Transfer pipe(duro lock) coupling boxup
28
A1280
manway box up
0.5d 01-May-16 09:36 PM
02-May-16 09:36 AM
29
A1290
Catalyst loading and Deblinding
1.0d 02-May-16 09:36 AM
03-May-16 09:36 AM
Catylyst unloading
Manhole Openning (for problem identification)
Opening of Catalyst transfer pipes of Rx1
Placement of chain pulley blocks for lifting center support plate
Opening of Center pipe support bolts of Rx1
provision of cutting of bolts(if not opened) of Rx1
TorqueTightening of center support pipe flange at 4.2 Nm of RX1
Placement of chain pulley blocks for lifting center support plate of RX2
Opening of Center pipe support bolts of RX2
provision of cutting of bolts(if not opened) of RX2
manway box up
Catalyst loading and Deblinding
Note: Activites highlighted as yellow are addtional activites in innovative approch
Actual Level of Effort
Remaining Work
Milestone
Actual Work
Critical Remaining Work
summary
Page 1 of 1
TASK filter: All Activities
© Oracle Corporation
Assessment of hydrogen unit revamp by heat exchange reformer
Prashant Parihar*, Ravi Kumar Voolapalli*, Pankaj Muley**, Nitin Jawale**, A. C.
Prabhune**, V. Suresh**
*Corporate Research and development Centre, Bharat Petroleum Corporation Limited
**Technical Services Department, Mumbai Refinery, Bharat Petroleum Corporation Limited
1. Introduction
BPCL Mumbai refinery (MR) has an installed crude processing capacity of 12 MMTPA. The
Refinery configuration has undergone several upgradation over the years. Various
secondary units have been added at MR. At present MR has the capability of producing BSIV grades of MS and HSD which are in agreement with the present Auto Fuel Policy. For
meeting future fuel quality norms it was identified to set up a Diesel Hydro treatment Unit
(DHT) of 2.6 MMTPA capacity along with associated facilities. In view of this, revamp of
existing Hydrogen Generation Unit was planned and currently being executed in MR to
meet the additional hydrogen requirement for DHT unit [1].
2. Capacity revamp of hydrogen unit by Heat Exchange Reformer (HTER)
Hydrogen Plant revamp for 30 % (approx.) increase in Hydrogen production is being carried
out by implementing Heat Exchange Reformer Technology of M/s Haldor Topsoe. Heat
exchange reforming as the name suggests, utilizes heat available in the process gas exiting
from Steam Methane Reformer (SMR) for steam reforming of additional feed. In a parallel
layout with SMR as depicted in figure 1, 20 - 30 % of the feed is split and taken to heat
exchanger type of reactor [2, 3, 4]. A proportion of heat is utilized in the process side which
reduces the steam production from the plant.
Schematic of hydrogen generation unit post revamp is depicted in figure 2. In this way
additional reforming capacity is realized by utilizing heat available in the existing unit that
was previously used for steam generation.
Figure 1: Typical layout of Hydrogen Unit revamp through a parallel HTER [3, 4]
The major advantage of this revamp scheme is increase in hydrogen production without any
increase of Fuel / Utility consumption.
3. Simulation, validation and optimization of revamp PFD’s
The main objective of the project was to target 125-130% hydrogen production after
revamping unit with HTER installation and operation on following feed stocks.

Design case 1: RLNG + 5500 Nm3/hr CCR PSA tail gas

Design case 2 : Heavy Aromatic Naphtha (HAN) + 5500 Nm3/hr CCR
PSA tail gas
Revamp flow sheets were established by licensor and shared with MR team for comments
and inputs. It was agreed to utilize previously developed steady state simulation model of
hydrogen generation unit for validating e a p ases PFD’s
li e so
provided. Using
e a p pfd’s; i -house simulation was incorporated with heat exchange reformer
configuration. All the major reactors (pre-reformer, reformer, heat exchange reformer, shift
reactors) were modeled in adiabatic manner in aspen plus software (Version 8.6).
Figure 2: Schematic of hydrogen unit post revamp
Process and flue gas heat recovery loop having all the exchangers and flash drums were
built in the simulation model. Combustion air and reformer fuel loop was integrated with
adiabatically modeled reformer, where in combustion of PSA offgas and fuel was providing
e ui ed dut fo adia ati
spe ifi atio s a d
efo
al ulato
e . Highlight of flo sheet si ulatio
lo ks i teg ati g
ei g use of desig
iti al pa a ete s of stea
efo
i g
process (S/C ratio, methane slip, reformer outlet temperature, combustion air flow rate,
oxygen content in flue gas, fuel requirement).
Validation of revamp case has helped to identify potential areas of energy savings across
the unit. First of its kind activity for team, utilizing simulation for identifying energy savings
in Reaction based (Gibbs Reactor) unit using Aspen Plus. Flow sheet simulation model
developed for entire plant for post revamp scenario was utilized to validate process
conditions provided by licensor (heat duties of exchangers, temperature and composition of
streams).
Du i g alidatio of PFD’s usi g si ulatio
odel it as oti ed that a e p o ess gas ai
cooler (133EA302) was recommended by licensor having 7.67 Gcal /hr duty as against
existing cooler having duty ~ 6.1 Gcal/hr as depicted in figure 3.
Figure 3: RLNG + CCR PSA Tail gas PFD (first issue) by licensor
Figure 4: Fi alized PFD o i corporatio of BPCL’s i puts for RLNG + CCR PSA Tail gas Case
Existing Process gas AFC was inadequate after revamp case for achieving the desired
thermal duty. Provision of additional bank was ruled out due to space constraint. Provision
of additional bank was ruled out due to space constraint thus the need was felt to optimize
additional heat duty in upstream heat exchanger network.
Thus the need was felt to optimize additional heat duty in upstream heat exchanger
network. It was identified using model that, additional heat available in the process gas can
be used to heat DM water. This resulted in net saving of 1.57 Gcal/hr of duty resulting in
avoidance of new process gas air cooler as recommended by licensor. Duty of DMW
preheater 133E308 was increased to 9.66 Gcal/hr from 8.14 Gcal/hr in base case PFD by
increasing DMW rate by 12 tph, this would also result in saving of @150 MT/D of LP steam
in the de-aerator due to preheating of additional quantity of DM water.
This suggestio has ee i o po ated i fi alized e a p PFD’s as depicted in figure 4.
Simulation output in terms of energy savings was communicated to licensor which was
ag eed a d i o po ated i P&ID’s. Sa i gs ealized
a oidi g e ui e e t of e p o ess
gas air cooler.
Economic analysis
Benefits have been estimated for operating with existing AFC over change recommended in
base case PFD provided by licensor. As per base case new process gas air cooler would have
consumed additional electricity equivalent to 1.57 Gcal/hr over existing process gas air
cooler at its maximum duty (6.1 Gcal/hr). Revamp case recommended by licensor has been
compared with pre-revamp configuration in table 1 for estimating benefits over
recommended case.
Component
Unit
AFC duty
POWER
Per month (30*24) hrs
Electricity rate in Mumbai
Cost of utility
Recurring saving/yr
New AFC as per
Existing AFC
Savings
base case of
at max duty
revamp
Gcal/hr
7.67
6.10
1.57
KW
8728
6941
1786
kwh/month
6283872
4997604 1286268
Rs/KWH
9.5
9.5
9.5
Rs. Crore/month
5.96
4.74
1.22
14.7 Crore/yr
Table 1: Recurring benefit analysis
4. Simulation of other revamp cases
Post validation of design cases, various feed scenarios were independently simulated using
in-house simulation model. Detailed process data was generated and unit capacities were
identified for various combinatio / o ditio s of feedsto k’s. Data generated is compared
with cases provided by licensor in figure 5. Simulation model successfully validated design
cases proposed by licensor.
Validatio a d fi alizatio of reva p PFD’s of design cases:
1.
RLNG + CCR PSA Tail gas
2.
HAN + CCR PSA Tail gas
3.
RLNG + CCR PSA Inlet gas
Check cases simulated in-house
1. RLNG only feed case
2. HAN only feed case
3. RLNG only feed case (operation at current capacity post revamp)
Table 1: RLNG + CCR PSA TAIL GAS
INPUT
Units
Design Simulation
NG Feed
kg/hr
20055
20055
CCR Offgas
kg/hr
3482
3482
Naphtha
kg/hr
0
0
S/C reformer
2.1
2.1
S/C heat exchange reformer
4
4
Methane slip from Reformer
5.5
5.5
Excess Air
1.1
OUTPUT
Stack O2 dry basis Mol fraction
0.018
0.020
H2 export kg/hr
7955
Steam to reformer kg/hr
55421
55423
Steam to exchange reformer kg/hr
9300
9294
Export Steam kg/hr
34632
34619
Temperature C
371
384
Pressure bar
41
41
Methane slip EX exchange reformer mole %
6.4
Net Methane slip mole %
5.7
5.6
PSA Offgas kg/hr
56973
57415
NG Fuel kg/hr
1945
1945
Combustion air kg/hr
184896
181183
Flue Gas kg/hr
243805
240365
Flue Gas Temperature C
173
175
Table 2: HAN+CCR Tail Gas
INPUT
Units
Design
Simulation
NG Feed
kg/hr
0
0
CCR Offgas
kg/hr
3482
3482
Naphtha
kg/hr
23104
23104
S/C reformer
2.1
2.1
S/C heat exchange reformer
4
4
Methane slip from Reformer
4.7
4.7
Excess Air
1.1
1.1
OUTPUT
Stack O2 dry basis Mol fraction
0.018
0.020
H2 export kg/hr
7192
7202
Steam to reformer kg/hr
63354
63353
Steam to exchange reformer kg/hr
8187
7999
Export Steam kg/hr
32310
31634
Temperature C
349
353
Pressure bar
41
Methane slip EX exchange reformer mole %
6.4
Net Methane slip mole %
5.0
5.0
PSA Offgas kg/hr
66721
66993.7
Naphtha fuel kg/hr
1731
1731
Combustion air kg/hr
177538
177538
Flue gas kg/hr
248014
248294
Flue gas temperature C
178
177
Table 3: RLNG +CCR PSA Inlet Case
INPUT
Units
Design
Simulation
NG Feed
kg/hr
21937
21937
CCR PSA Inlet
kg/hr
750
750
Naphtha
kg/hr
0
0
S/C reformer
2.1
2.1
S/C heat exchange reformer
4
4
Methane slip from Reformer
5.6
5.6
Excess Air
1.1
1.1
OUTPUT
Stack O2 dry basis Mol fraction
0.018
0.020
H2 export kg/hr
7756
7745
Steam to reformer kg/hr
52489
52489
Steam to exchange reformer kg/hr
9013
9013
Export Steam
32318
32807
Temperature C
365
369
Pressure bar
41
Methane slip EX exchange reformer mole %
6.3
Net Methane slip mole %
5.9
5.7
PSA Offgas kg/hr
53665
54162
NG Fuel kg/hr
1579
1579
Combustion air kg/hr
175472
175472
Flue Gas kg/hr
230715
231213
Flue Gas Temperature C
172
174
Figure 5: Typical simulation output generated for revamp cases
5. Comparing base case with revamp
In-house model has been utilized successfully for comparing current configuration versus
revamp case at current design capacity of 148 TPD. Comparison highlighted that revamp
case is beneficial even if it is operated at current capacity; saving to the tune of Rs. 6.3
Crores (while accounting for reduction in steam generation post revamp) can be realized
through reduction in operating cost post revamp. Detailed calculations are elaborated in
table 2.
Existing Unit Design @148.5 MT/D
UOM
Hydrogen production
Revamp case with HTER @ 189 MT/D
only RLNG as feed
MTD
only RLNG as feed
1.00
1.00
Particulars
Qty
Rate/MT
Rs./lakhs Rs./MT
Qty
Rate/MT Rs./lakhs Rs./MT
Of H2
Of H2
RLNG Feed
MTD
3
56,101
1.68
168,303
2.90
56,101
1.63
162,693
RLNG Fuel
MTD
0.45
56,101
0.25
25,245
0.33
56,101
0.19
18,513
MTD
(6.4)
3,990
(0.26)
(25,536)
(3.60)
3,990
(0.14)
(14,364)
HP steam export actual (Negating steam to Jackets @
150 MT/D)
Sea Cooling water consumption (Rate is Rs./TMT)
TMTD
-
Water
MTD
Assuming no considerable
change
Assuming no considerable
change (although DM water
consumption has reduced).
Net boiler feed water/DM water consumption
MTD
Compressed Air
000nm3
-
-
Power consumption (Rate is Rs. Per Kwh)
KWH
-
-
Operating cost at NHGU,( per MT of H2 production)
Rs/MT
Difference in operating cost , existing - revamp , (A-B)
Rs/MT
A
-
-
168,012
B
1,170
Saving in Operating cost for 148 MT/D of H2 production
Rs in Lakhs
Saving in Operating cost per year (365 days)
Rs in Crores
1.73
6.32
Table 2: Comparing Heat Exchange reforming revamp with current configuration
6. Conclusion

Achieved operational excellence while finalizing revamp case by using state of
the art simulation platform fo
alidatio of e a p pfd’s p o ided
li e so .
BPCL team optimized revamp configuration and identified potential areas of
-
166842.2
energy savings across the hydrogen unit. Space and time constraints overcome,
allowing timely execution of revamp configuration.

Identified heat integration opportunity upstream of Process Gas Air Cooler.
Study resulted in establishing adequacy of existing air cooler and avoided need
of new facility which was recommended by licensor. Revamp case finalized after
systematic validation and incorporating energy efficient configuration. Capex
saved as new process gas air cooler recommended by licensor was avoided by
implementing additional heat exchange between process gas and DMW
(recurring energy savings 1.57 Gcal/hr ~ 150 TPD LP steam saving in deaerator).
Capital savings to the tune of Rs. 0.5 Crores were achieved by avoiding new
process gas air cooler recommended by licensor. Savings in operating cost to the
tune of Rs. 14.6 crores per annum were realized by avoiding 1.57 Gcal/hr duty in
process gas air cooler.

Post validation of base case, various feed scenarios were independently
simulated using in-house simulation model, resulting in one time savings in
consultancy charges of licensor.

Simulation model ready for monitoring and optimization of hydrogen unit with
heat exchange reformer configuration post start up in May-2017.
References
1. Feasibility Report on Diesel Hydro-Treater (DHT) Unit and Associated Facilities, January
2015, Bharat Petroleum Corporation Limited.
2. Ib Dybkjaer, Sandra Winter Madsen, Niels Udengaard, Revamp options to increase
hydrogen production , Petroleum Technology Quarterly, 2000, Q2, 93-97.
3. Henrik Olsson, Poul Rudbeck a d Ki
Capa it
E io
Heat E ha ge Refo
H. A de se , Addi g H d oge
i g , XIV Refi e
e t Challe ges fo the H d o a o
P odu tio
Te h olog Meet RTM o ’E e g &
Se to ’ 20-22 September 2007, Kovalam,
Trivandrum, India
4. Jack Heseler Carstensen, Additional hydrogen production by heat exchange steam
reforming , Pet oleu
Te h olog Qua te l , 2010, Q4, 47-51.
IMPLEMENTATION OF RELIABILITY CENTERED
MAINTENANCE (RCM) AND FAILURE MODE AND
EFFECTS ANALYSIS (FMEA) IN ESSAR OIL
Jayanti Vagdoda, GM, Head Rotary, Reliability and ISBL-2, Essar Oil
Aditya Trivedi, Sr. Manager-Maintenance, Essar Oil
υ
ABSTRACT
In the wake of the emerging challenges and a desire to achieve and maintain pacesetter
performance, it becomes imperative to identify tools to keep the units running with optimized
expenditure. Lot of operating and maintenance data, which is very valuable, is generated for
each asset. This data is stored in various point solution like SAP-PM, process historian, MS-Excel
sheets, different portals, etc.
To achieve high reliability of the assets and ensure that they continue to function as desired, it
is necessary to analyze this data to obtain a holistic picture of the asset and develop a
comprehensive maintenance strategy of the asset.
We have Reliability Centered Maintenance (RCM), Failure Mode and Effects Analysis (FMEA)
and strategy optimization methodologies available which can give us such comprehensive
maintenance strategies. To use these methodologies we require this operation and
maintenance data of the asset to be in a single place as referring to this data from various point
solutions is a cumbersome process.
We have used Meridium tool to collate this data into one place and use the above mentioned
strategy building methodologies. Before we start collating the data we carried out some
changes in SAP-PM module regarding equipment taxonomy, catalog profile and notification
types to standardize and ease data handling.
We have also revised asset risk matrix and based on criticality of assets from the revised
criticality matrix, we selected the suitable methodology for strategy creation.
These methodologies being new for us, we decided to go for phase wise implementation with
38 critical and unique systems from the entire refinery, to be taken up as phase-1. Phase-1 was
done under guidance from subject matter experts and our personnel were trained. Phase-2 we
have planned to do in-house.
We have found improvements in existing maintenance strategy of some of the phase-1 assets
to increase their reliability.
To conclude we can say that these methodologies provide an exhaustive maintenance strategy
for an asset or a system. It enables the personnel operating and maintaining the asset to work
harmoniously for maintaining the asset reliability.
φ
INTRODUCTION
It is a fast changing world with new technologies being invented and applied on a continuous
basis. There is a plentiful of emerging challenges to be faced by the industry. So, in a desire to
achieve and maintain pacesetter performance, it becomes imperative for the organizations to
think out-of-the-box and find out a fresh way of doing business as compared to the old
fashioned way.
The installed assets are very valuable resource for a refinery. It is very much essential that this
resource is maintained in a condition to enable it to function as per its desired performance. It
is also essential to do this at an optimized cost.
There is a lot of valuable operation and maintenance data generated continuously for each
asset. There are various solutions viz. SAP, process historian, MS-Excel sheets or various web
portals, where this data is getting generated/stored. A tool is, therefore, required to collate this
data to give a holistic view of the asset and facilitate analysis of this data for generating its
maintenance strategy. A maintenance strategy developed this way empowers the maintenance
department to maintain the asset reliability at optimized cost.
SAP-PM module changes before collating the data
SAP-PM is a solution used for capturing work history of an asset. Before collating this data with
other solutions, it is desired to have SAP-PM standardized to provide adequate granularity to
the data captured as well as facilitate fast entry of data. This granularity will ease handling and
sorting of this data. We have done the following changes to achieve data granularity and ease
data entry:
1. Equipment Taxonomy Mapping: Ta o o
ea s lassifi atio . We have odified
O je t T pe field of SAP e uip e t aste data to contain the details of every asset
based on its construction. Few examples are given below:
2. Catalog profile cleansing: A review of the existing catalog profile was done to reduce
the items listed in maintainable item, damage code and cause codes fields. This makes
the data crisp and enables fast entry of work history.
χ
3. Notification type reduction: There were more than
required types of notifications in SAP which created
confusion in selecting correct type of notification while
creating notification. To ease this process the
notification type not used in SAP were disabled for
creation.
Collating the data
We have used a software tool supplied by M/s Meridium for
collating the data. The tool integrates with SAP-PM, Process
historian and our web portals to bring all the data on a single platform. It can also take data
inputs from MS-excel worksheets. The collated data can then be viewed against that asset, on a
single page. Queries can also be made to search the desired information from the database
and present it in desired format.
Once we get the holistic picture of an asset, risk based approaches like RCM and FMEA along
with strategy optimization can then be used to formulate exhaustive maintenance strategies to
achieve highest reliability at optimized cost.
Risk Matrix to decide a suitable maintenance optimization technique
We have developed an asset risk matrix to decide a suitable maintenance strategy optimization
technique. The risk matrix of an asset defines all risks associated with Safety, Environment,
Asset loss, Business loss and reputation for a worst case scenario.
Reliability Centered Maintenance (RCM)
Reliability Centered Maintenance: A process used to determine what must be done to ensure
that any physical asset continues to do what its users want it to do in its present operating
context.
RCM-II, by John Moubray
ψ
The RCM process is defined by asking seven questions about the asset or the system under
review:
1.
2.
3.
4.
5.
6.
7.
What are the functions of the asset in its present operating context? (Functions)
In what ways does it fail to fulfill its functions? (Functional Failure)
What causes each functional failure? (Failure Mode)
What will happen when each failure happens? (Failure Effect)
In what ways does each failure matter? (Failure Consequences)
What should be done to predict or prevent each failure? (Recommendations)
What should be done if a suitable task is not found? (Default Action)
RCM is done for super-critical assets as these are key for the business and hence maintaining
their reliability if of prime importance.
Failure Mode and Effects Analysis (FMEA)
FMEA focuses on the asset and is an asset based approach to formulate maintenance strategy.
This is used for critical assets.
The FMEA process asks the following questions to the analysis team:
1.
2.
3.
4.
5.
In what ways the asset can fail? (Failure mode)
What will happen when each failure happens? (Failure Effect)
In what ways does each failure matter? (Failure Consequences)
What should be done to predict or prevent each failure? (Recommendations)
What should be done if a suitable task is not found? (Default Action)
These questions are answered for each component of an asset.
Strategy Optimization
Strategy optimization process is comparing the existing strategy (actions) with the proposed
ones, in terms of balancing the risks mitigated and the cost of each action. This methodology
can be directly applied to Semi-critical or non-critical assets and is used for optimization of
strategies proposed by RCM or FMEA.
Implementation
RCM and FMEA are exhaustive studies of the assets and require a multidisciplinary team
comprising of experts of the asset or system, under study, from Operations, maintenance,
rotary, reliability and process departments. This team then sits together and does
ω
brainstorming and knowledge sharing to get holistic picture of a particular asset or system
which is used to develop the maintenance strategy for that asset or system.
Being a new approach towards maintenance, we have gone for a phase wise implementation in
our refinery. Phase-1 comprised of 38 critical and unique systems selected from the refinery
units. The studies were conducted in presence of subject matter Experts to handhold and guide
us. Phase-2 is the application of these methodologies to the remaining systems with our
internal expertise and sample formats developed during phase-1.
A dedicated core team was formed for developing the software and completing the phase-1
studies. The following process was followed:
1. Business Blue Printing: Process of identifying changes required in the standard software
to suit it as per our requirements. This is a joint exercise between us and the software
provider.
2. Configuration: Changing the standard software to suit our requirements as per blue
printing process. This activity is done by the software provider.
3. Deployment and User Acceptance Testing: Setting up the client servers and testing of
the software by client users to ascertain that the configured software meets the client
demands. This is a joint exercise between us and the software provider.
4. User Training: Training of our users on the software and associated methodologies.
Training is imparted by Subject Matter Experts (SMEs) deployed by software provider.
5. Consulting or Studies: Coordinating with designated SMEs, deployed by software
provider, and conducting sessions at our site for carrying out RCM and FMEA analyses of
phase-1 systems.
Some findings from the phase-1 studies
We found may new recommendations from the studies that can be implemented to improve
the reliability of the assets. Some of these are listed below:
1. Cooling air provision in Utility-2 air blower to increase bearing life.
2. Flushing line size increase in process centrifuge of utility-2 to increase screw life.
3. Online vibration measurement system installation in VR pumps to proactively identify
and prevent major failures.
ϊ
Conclusion
These methodologies provide an opportunity to all personnel, connected with the asset, to
come together and do brain storming to have all encompassing maintenance strategy for the
asset. They get a new insight of their role in preserving the function of the asset and then they
work harmoniously to keep the asset functioning at per its desired performance.
References:
1. Moubray, John, RCM-II, Second Edition, Butterworth-Heinemann.
ϋ
21st Refinery Technology Meet
Upgradation of Pitch from Slurry phase Resid Hydrocracking
Bhavesh Sharma, Kanuparthy Naga Raja, Peddy V C Rao, G Sriganesh
HP Green R&D Centre, Hindustan Petroleum Corporation Limited, KIADB Industrial Area,
Devanagundi, Hoskote, Bangalore, Karnataka, 560067, India
Abstract:
In Slurry Phase Resid Hydrocracking, pitch forms a portion of the product (bottoms) which
has a boiling point above 540° C. The utilization of Pitch is limited due to its high viscosity
and high levels of undesired components which makes it a very low value product. One of
the ways to put this pitch to use is to gasify it to form Synthesis Gas for power generation or
Fuels and Chemicals. This study presents a techno-economic analysis of Pitch Gasification
as a part of the Slurry Hydrocracking Process. A model has been developed in Aspen Plus
to study the gasification of pitch obtained from Slurry Hydrocracking Process to understand
the thermodynamic feasibility of the said process. Furthermore, a comparative study with the
gasification process of different kinds of coal has also been made as a part of this work in
order to have an understanding of the economic benefits derived by gasification of
pitch.Thus an integrated complex of Slurry Hydrocracker with Co-Gasification of pitch to
upgrade the refinery bottom products is investigated.
Introduction:
With the worldwide increase in demand for high value petroleum products such as middle
distillates, gasoline and lube, refiners are focused on the maximization of yield of liquid
products by various processes. Upgradation of refinery bottoms such as vacuum residue
which is the bottom product from vacuum distillation unit and is cheaply available in huge
quantities has sparked an interest in this regard. Hydrogen addition processes for
upgradation of heavy oil or vacuum residue are preferable to carbon rejection processes
owing to the fact that the former produce more distillate liquid products with high H/C
ratios.[1] Hydrocracking is one such hydrogen addition process which has proved to be more
reliable in achieving higher conversions and handling more difficult feeds, when carried out
in slurry phase reactors.
Slurry phase hydrocracking is marked by the introduction of a slurry of heavy oil with the
catalyst into a high temperature and high pressure reactor with Hydrogen. Gas, naphtha,
diesel, vacuum gas oil and the heavy unconverted portion form the main products of
hydrocracking of vacuum residue in slurry phase.Pitch forms a portion of the product
21st Refinery Technology Meet
Visakhapatnam
(bottoms) which has a boiling point above 540° C and cannot be readily used.One of the
ways to put this pitch to use is to gasify it to form Synthesis Gas for power generation or
Fuels and Chemicals. A typical pitch obtained as a result of slurry hydrocracking of vacuum
residue derived from Basrah crude as feed contains 4.94% Sulphur, 0.13% Ash content and
has a heating value of 38.889 MJ/Kg. This pitch can have a chemical value in terms of
heating, similar to that of coal, which is subjected to gasification and is useful as a precursor
for chemicals or in power generation.
Gasification Technology provides the advantage of producing syngas which can be used for
both power generation and chemicals production. The gasifier feedstock include coal,
natural gas, biomass wastes, petcoke and refinery residues[2]. While a conventional coal
gasification based power plant has a typical plant efficiency of about 35%, an integrated
gasification combined cycle (IGCC) power plant which includes both gas and steam turbines
can have a plant efficiency of about 45%. Even though the latter involves a higher capital
cost, it is favoured over the former on account of its higher efficiency and plant economics.
This study aims at understanding the thermodynamic feasibility of the gasification of pitch
obtained during slurry hydrocracking of vacuum residue and compare against conventional
coal based IGCC process. An integrated approach thus can be envisaged when Pitch, the
bottom product of Slurry Hydrocracking is to be utilized effectively.
Modelling Strategy Adopted:
Model of a typical Integrated gasification combined cycle has been developed using Aspen
Plus. Feeds were defined with the help of user properties and custom modules.
Figure 1: Modelled Integrated Gasification Combined Cycle Scheme
In the developed model, a pitch/coal-water slurry reacts with oxygen in the gasification unit.
The solid products are removed as ash from the bottom while the gaseous products are
cooled and sent for removal of ammonia and hydrogen sulphide. The resulting clean gas
stream is then passed on to the combustion chamber following which power is generated by
21st Refinery Technology Meet
Visakhapatnam
the gas turbine. The turbine exhaust further passes through the Heat Recovery Steam
Generator to generate more electricity.[3]
2500 tons per day of pitch/coal was taken as a basis for developing the said model.
Pitch/coal was modelled as non-conventional solid in Aspen Plus, the inputs being proximate
analysis, ultimate analysis and heating value.[4] Gasifier was operated at a pressure of 450
psig while the heat duty from the yield reactor, which decomposed the pitch/coal into its
constituent elements, was carried with energy stream to the Gasifier. The gasification
temperature is maintained between 800 0C and 1200 0C depending upon the type of feeds
involved in the study.
Base case was run using Pitch produced as a result of slurry hydrocracking as the feed for
gasification. Subsequently, simulations were done using the developed model with three
different kinds of coal, namely; Indian Coal (Chandrapur), Australian Coal (Moreton Basin)
and Texas Coal(Bowie). Furthermore, 50% Pitch and 50% of coal were also used for
simulations in order to understand the feasibility of Pitch gasification.
Results and Discussions:
Since Pitch and Coal were modelled as non-conventional solids in Aspen Plus, certain
properties of the feed needed to be specified for developing the model. Table 1 tabulates the
properties of different feeds used for gasification in the model.
PROPERTIES OF FEED
Moisture,%
Fixed Carbon,%
Ash,%
Volatile Matter,%
Carbon,%
Hydrogen,%
Nitrogen,%
Sulphur,%
Oxygen,%
Higher Heating
Value,MJ/Kg
0
66.46
0.13
33.41
84.49
6.54
0.31
4.94
3.59
4.5
27.5
47.5
20.5
41.89
2.66
1.27
0.9
5.78
AUSTRALIAN
COAL[5]
11.1
41.7
15
32.2
69.8
6
1.5
0.7
7
38.889
15.2711
25.24
PITCH
INDIAN COAL
TEXAS COAL[6]
24.121
11.97
26.39
6.82
54.82
59.83
6.11
1.7
0.74
24.8
Table 1: Proximate and Ultimate Analysis of different feeds used for modelling
For understanding the thermodynamic feasibility of using Pitch for Gasification, Cold Gas
efficiencies were calculated for each of the cases of feedstock. The cold gas efficiency was
calculated as ratio of higher heating value of fuel gas produced to the higher heating value of
feed. Net Power Production was calculated as the total net power generated by the HRSG
and the Gas Turbine minus the power required to operate the Air Compressor and Oxygen
Compressor in the Simulation. The results are tabulated in Table 2.
21st Refinery Technology Meet
Visakhapatnam
FEED
Cold Gas Efficiency,%
Net Power Production, MW
Pitch
Indian Coal
Australian Coal
Texas Coal
86.05
66.39
76.38
75.14
364
118
222
204
Table 2: Net Power Production and Cycle Efficiency for Pitch and 3 different Coals
From Table 2, it is evident that Gasification of Slurry Hydrocracking Pitch is a feasible option
which led to the highest cold gas efficiency of 86.05% amongst all the feedstock fed to the
gasifier. It was also seen that Indian Coal which has the Highest Ash content had the poorest
efficiency. Furthermore, net power generation by the combined cycle with pitch as feedstock is
far greater in comparison to that with the other feedstock.
Apart from using the three different coals and the Pitch for simulation, a different case of 50%
Pitch and 50% coal was run for all the three types of coals. Results for the same are tabulated
in Table 3.
FEED
50% Pitch+ 50% Indian Coal
50% Pitch + 50% Australian Coal
50% Pitch + 50% Texas Coal
Cold Gas Efficiency,%
70.58
81.43
80.03
Net Power Production, MW
223
288
273
Table 3: Net Power Production and Cycle Efficiency using 50-50% combination of Pitch and Coal
It can be further noted that a combination of pitch along with coal can lead to an improvement
in cold gas efficiencies and net power generation of the respective coals.
According to the USDOE report[8], the capital cost of a single unit coal-based IGCC without
CCS is 4400$/KW. The capital cost of Pitch-based IGCC model that was developed using
Aspen Plus is estimated to be 4421 $/KW, which is in close agreement with the former.
However, as the case with Gasifier, handling of coal admixed with pitch and the resultant
product ash remains a concern and need to be dealt with through engineering excellence.
Conclusions:
The gasification of Pitch obtained as a result of slurry hydrocracking has been modelled using
Aspen Plus. It has been observed that a cold gas efficiency of 86% is obtained with Pitch
gasification. Pitch which has a higher heating value and negligible ash content as compared to
coal led to the highest efficiency as well as Power production by the combined cycle. It is also
observed that when a combination of Pitch and Coal is fed to the gasifier, an improvement in
21st Refinery Technology Meet
Visakhapatnam
the power generation and cold gas efficiency is observed in comparison to the standalone coal
gasifiers. The results provide an insight into the possibility of co-gasification of Pitch along with
Coal which can be used for chemicals or power generation. This study thus is limited to
thermodynamic feasibility and paves way for the detailed investigation of an integrated
complex of Slurry Hydrocracking with Gasification facilities.
References:
[1] J.G. Speight, The Desulfurization of Heavy Oils and Residua, Marcel Dekker Inc., New
York, 2000.
[2] Andrew J. Minchener, Coal gasification for Advanced Power Generation, Fuel 84(2005),
2222-2235
[3] Zachary Hoffman, Simulation and Economic Evaluation of Coal Gasification with SETS
Reforming Process for Power Production, Louisiana State University, May 2005
[4] Aspen Plus. “Getting Started Modelling Processes with Solids.” Aspen Technology, Inc.
Burlington ,MA. 2013.
[5]Dr. Chris Spero, Utilization of Queensland’s low and high volatile coals in power
generation, EU-Australian Coal Conference, Germany, 24-25 Sept 2001.
[6] W.L.Fisher, Lignites of the Texas Gulf Coastal Plain, Bureau of Economic Geology,The
University of Texas, October 1963
[7] Ola Maurstad, An Overview of Coal based Integrated Gasification Combined Cycle
(IGCC) Technology, Massachusetts Institute of Technology, Sep 2005
[8] Updated Capital Cost Estimates for Utility Scale Electricity Generating Plants, U.S.
Department of Energy Washington, April 2013
[9] Neville A.H. Holt,Integrated Gasification Combined Cycle Power Plants, "Encyclopedia of
Physical Science and Technology", September 2001
21st Refinery Technology Meet
Visakhapatnam
Deployment of Parallel Solve
Automation in LP models for speedier
Spot Crude oil evaluation
21st Refinery Technology Meet
20th to 22nd April 2017, Vishakhapatnam
Submitted by:
Mridusmita Goswami
R Jerold
Kabidas Mandal
S N Pandey
Asstt Manager (Optimization), IOCL, Corporate Office, New Delhi
Manager (Optimization), IOCL, Corporate Office, New Delhi
General Manager (Optimization), IOCL, Corporate Office, New Delhi
Executive Director (Optimization), IOCL, Corporate Office, New Delhi
1
1 Introduction
Indian Oil is India’s flagship Maharatna national oil company with business
interests spanning the entire hydrocarbon value chain – from refining, pipeline
transportation and marketing of petroleum products to Research & Development,
Exploration & Production, marketing of natural gas and petrochemicals. Indian Oil
Group (including two refineries of its subsidiary company Chennai Petroleum
Corporation Ltd. (CPCL)) owns and operates 11 of India’s 23 refineries with largest and
most extensive network of retail outlets with more than 24000 touch points consisting of
depots, terminals, LPG bottling plants, AF stations etc.
Indian Oil’s varied business interests which range from crude procurement to
product distribution and product logistics requires a robust, effective and integrated
supply chain management which takes into account the intricacies and complexities of
the entire system. The large number of refineries with varied complexities also
necessitates different crude and product slates. It utilizes LP tool, Refinery &
Petrochemical modelling system (RPMS) for its planning purposes. The integrated
planning model thus enables the refiner to provide the best supply chain solution
keeping in view the available constraints. It thus enables to maximise margins in the
system.
2 Integrated Planning
The Integrated Planning Model (IP) is a combination of detailed refinery LP
models of Barauni, Gujarat, Haldia, Paradeep, Mathura, Panipat Refinery &
petrochemical, Guwahati, Digboi & Bongaigaon, refineries along with simple models
(Extreme point) of CPCL & CBR refineries. The model also captures the cross country
pipelines, other logistic networks for all the products, connecting all depots and
terminals with different modes of transportation across India. The IP model thus
captures a huge magnitude of around 85000 variables and 35000 constraints capturing
all of its refineries, pipelines and marketing structures.
The IP model utilizes the inter relations among crude procurement, refining
operations and product distribution to generate globally optimal, market feasible
solutions. The following figure illustrates the data flow in the IP model. Thus, best
refinery production planning and product supply and distribution planning are provided
by the IP model.
2
Fig 1: Data flow in Integrated Planning (IP) model
The Integrated Planning model has the following inputs:
1) Crude availability at ports
2) Grade wise products Demand at locations
3) Committed Exports/Imports / Exchanges with Other Marketing Companies
(OMCs)
4) Planned refineries process units shutdown schedule
5) Seasonal variation in products specifications
6) Crude landed cost at ports
7) Products prices at demand location
8) Desired Inventory build-up / depletion at depots level
The IP model provides the refiner with the following outputs:
1) Refinery wise throughput and crude allocation
2) Crude requirement
3) Refinery wise product slate
4) Purchases, exports, imports, inventory build up/depletion
The integrated planning model has thus enabled to take decisions spanning the
entire hydrocarbon chain and has also yielded in improvement in refining margins.
3
3 Crude Evaluation
The inputs given to the IP model and the output received from it in the form of a
production plan formulates the ground-work for the evaluation of spot crude for a tender
based procurement process. The crude evaluation process involves the
combination/grades of crudes be considered for processing in the IP model one by one
and calculation of the corporate objective functions. In a spot-crude tender evaluation,
crude evaluation method involves indifference value calculation of each crude
combination/grades using linear programming and selecting the crude
combination/grade which gives the maximum corporate objective function. The objective
function is defined as difference between sales realization (domestic demand, exports
etc.) and the cost incurred in all of the inputs (crude oil, catalysts, utilities etc.).
Obj. Function = Σ Sales realization (Domestic demand + Exports +
Exchanges) - Σ (Feed cost at ports + Products imports/Purchases + Logistic cost
for crude and products + Catalyst & Chemical cost + Utilities purchase)
Crude Evaluation : Schematic
Crude 1
OBJ. FNC. 1
H2
Crude 2
OBJ. FNC. 2
H1
Crude 3
OBJ. FNC. 3
 Output is Corporate Objective Function i.e. Margin of all the three divisions
 Refineries
 Marketing
 Pipelines
1
Fig 2: Crude evaluation process
4
H3
Indian Oil’s crude oil basket consists of crudes ranging from heavy crudes of
around 20 API to condensates of more than 55 API. Indian Oil processes different
crudes of varied ratios to mitigate the extreme properties and also to widen the crude oil
basket, which results in large number of combinations/grades. Owing to the huge
capacity of the refineries, the quantity of crude requirement is enormous. The cost of
crude oil is the major contributor of operating costs; hence continuous efforts are being
made to reduce the crude oil input costs to the refineries.
In the year 2012-13 around 250 combinations/grades of crudes were evaluated
for spot crude oil purchase. With inclusion of higher number of opportunity crudes and
continuing efforts to widen the crude oil basket, the number of combinations/grades to
be evaluated for each spot crude oil purchase has increased to over 800. An LP model,
with high cycle- time, posed a hard constraint by limiting number of cargoes purchased
per process (tender), thereby necessitating a repetition of the tender process for
subsequent purchases. Also with limited provision of flexibility to increase the volume
exposure according to varying market sentiments, this becomes a Herculean task as
the tender evaluation process is to be carried out around 3-4 times/month and the total
cycle time taken for an evaluation is nearly 2 days with an average per case run time of
around 15-20 mins.
4 Developments in crude evaluation process
Indian Oil envisaged reducing the run time for spot crude evaluation in order to
take advantage of dynamic market scenario. This will enable to utilize the huge potential
in tapping the crude-oil market opportunities and maximizing exposure to procurement
of high-valued cargoes. Hence, it embarked on a journey to improve on its LP modelling
techniques and put into use the latest development in technology and also in its LP
Solver. Extensive research was carried on at the in house level to improve upon the run
time of cases. Regular cleaning and proper maintenance of the LP models of refineries
and IP model was carried out. Also exercises related to reports consolidation yielded
favourable results. All of these efforts have resulted in the reduction of run time up to
the extent of 60%.
The linear programming tool, RPMS was upgraded and a newer version with
latest developments in modelling and solver was released by Honeywell. This further
enabled us to dig further and come with more solutions to minimise the run time for
crude evaluation process. Inputs to the LP models which results in matrix formation for
the solve process was consolidated and simplified in terms of size and structures. This
resulted in a further time reduction of 60%, which is around four mins/case.
5
5 Parallel Solve Automation
Indian Oil Corporation dedicated the year 2017 as the year of “Innovation and
Technology”. Keeping in line with the company’s vision, we visualized a concept of
parallel solving and along with M/s Honeywell Automation India Limited (HAIL) devised
a novel automation method for crude evaluation. This method combines both the
hardware capabilities like processor capabilities, multi-core, RAM etc and software
capabilities like latest LP solver, development in modelling techniques etc. and executes
parallel solving technique. The schematic diagram of the process is been illustrated in
the following figure 3.
Virtual PC-i
Virtual PC-n
PC-1
Virtual PC-i
Virtual PC-n
PC-2
PC-A
Virtual PC-i
Virtual PC-n
PC-n
Fig 3: Parallel solve process
6
The parallel solve process puts into use the processor frequency, core into use and
initiates multiple parallel solve processes in the user PCs. This process has been
developed in such a way that the automation tool takes into account and performs the
entire process from start to end viz. reports consolidation thereby reducing the manual
interference in the processes; hence minimising chances of any error. The automation
process is further assigned to work in an automated Multi-PC environment, which
utilizes other hardware resources. Each of the PCs further divides the job assigned
based on the number of hardware resources like virtual PCs. The user PC takes into
control the other PCs and performs the assigned task of crude evaluation. Thus only
one user can utilize the hardware and software capabilities of different user systems
and put them into effective use for crude evaluation purpose and complete the target in
a shorter time.
The development in the crude evaluation process by deployment of the parallel
solve technique has further helped in reduction of run time by around 35%. The entire
process of evaluation of around 800 cases for crude evaluation can now be performed
in a time frame of 3 hours.
6 Path Forward
The automation technique has led to the evolution of conventional crude
evaluation process and has paved the way for utilising hardware resources by parallel
solve techniques, use of available hardware resources. It has empowered us in quick
and precise decision making, to take advantage of changing crude oil markets. It has
also enabled the employment of human resources in other avenues for margin
improvement. This method has enabled to save time and has opened newer avenues of
profit making and has brought about dynamism in the crude procurement processes.
Keeping into consideration the benefits the automation has given to the
corporation, it has been envisaged to move forward and employ high end servers for the
said purpose. The use of servers with better hardware configurations will further enable
us to reduce the time taken for crude evaluation of around 800 cases to be completed
within an hour. Also, it has been acknowledged by M/s Honeywell Automation India
Limited (HAIL) to develop more on their LP modelling techniques and the use of latest
solver. This will also enable in enhancing the crude oil basket by inclusion of newer
grades/combinations of crudes, taking advantage of dynamic market scenario and
further improve refining margins.
7
7 Conclusion
This paper demonstrates how Indian Oil has taken advantage of the
developments in technology in its evolution of crude evaluation techniques and put them
into use for its crude evaluation process. Extensive exercises of LP model cleaning and
maintenance has been glued together with development of hardware resources and
software capabilities. The journey of time reduction in crude evaluation process that was
embarked upon has yielded fruitful results and have opened up newer pastures of
sequential ranking, real time decision making, and purchase of crudes through trading
desk to name a few. All these efforts have resulted in higher refining margins and
maximising profits.
8
INDIANOIL CORPORATION LIMITED,
GUWAHATI REFINERY, NOONMATI,
GUWAHATI,
ASSAM - 781020
Utilization of low cost Coal/Petcoke
as fuel for power generation in
place of fuel oil
GUWAHATI REFINERY
Utilization of low cost Coal/Petcoke as fuel for power generation in place of
fuel oil
INDIAN OIL CORPORATION LIMITED
(REFINERIES DIVISION)
GUWAHATI REFINERY
NOONMATI: GUWAHATI -20
1. INTRODUCTION :
Thermal Power station of Guwahati refinery consumes 6.5 to 7.0% of fuel on crude
t’put which is the highest amongst all IOCL refineries because power demand has been
catered by Steam turbines and for that fuel oil and gas are fired in Boilers. Thermal
Power station of Guwahati refinery consists of five oil & gas fired boilers and three
Steam turbines which are of extraction cum condensing type.
Guwahati refinery Thermal Power station consists of the following boilers :
Installed
Operating
Operating
Yr. of
capacity (MCR),
pressure,
Temperature,
Make
commissioning
TPH
Kg/cm2-g
deg C
39
450
Boiler 3
20
Rumanian
1962-64
39
450
Boiler 4
20
Rumanian
1962-64
39
450
Boiler 5
40
M/s IJT
1994
39
450
Boiler 6
50
M/s Thermax
2004
39
450
Boiler 7
50
M/s Thermax
2004
The two Rumanian Boilers (Boiler 3 and Boiler 4) have been in operation for the past 50 years and
Boiler
have outlived their services and are operating at low efficiency of 71 % – 74 % (against design
efficiency of 89 %) with a maximum capacity of 15-16 TPH against MCR of 20 TPH. Also, spares of
these boilers are not available for carrying out proper maintenance and RLA study of these boilers
carried out by external agency, M/s Energo Engineering Projects Ltd, has declared that these boilers
must be condemned.
PROPOSAL FOR INSTALLATION OF PETCOKE BOILER:
The opportunity of utilizing Petcoke generated in DCU, as fuel for boilers for steam generation was
explored and it was found that not only will it meet the future steam demand of the refinery, but also
provide significant GRM benefit, since the price of Petcoke is about 1/3rd the price of Fuel oil.
Technology adopted :
Utilization of low cost Coal/Petcoke as fuel for power generation in place of
fuel oil
CFBC technology (Continuous Fluidized Bed Combustion) is the most suitable choice for firing
Petcoke due to its ability to maintain same pollution standards as an Oil & Gas fired boiler. The
advantage of processing low sulfur Assam crude at Guwahati refinery results in comparatively
manageable 0.7 wt % sulfur content in Petcoke.
a) PET COKE PRODUCTION AT GUWAHATI REFINERY
The Delayed coking process is a thermal cracking process for upgrading petroleum residues into
lighter gaseous and liquid products and solid coke. The petroleum residue feedstock is heated in a
specially designed heater to high temperature with a brief residence time. The thermal cracking
reactions start in the heater coils and are completed in coke drums. The solid product (Green Coke) is
retained in the coke drums and recovered through decoking operation. There are three types of coke
namely fuel grade, anode grade and electrode grade are produced from delayed coking process.
Delayed Coking Unit at Guwahati refinery was commissioned in April 1962 with the processing
capacity of 0.33 MMTPA of RCO. In 1997, with the intent of improving the energy consumption and
maximizing the yield of middle distillates and increasing the RCO throughput to 0.44 MMTPA,
Technip KT India Ltd. (TPKTI) formerly Kinetics Technology India Limited was selected to revamp
the unit and provide the Basic and Detailed Engineering, Procurement, Construction and Project
Management services for the project. The DCU unit was successfully commissioned in May 2000 after
revamp job.
Guwahati produces Anode grade Petcoke having typical moisture content of 6-8 wt%,
ash content of 0.2 – 0.3 wt% and sulphur content of 0.6 – 0.7 wt%. Typical production of Pet coke ex
Guwahati DCU is 14-15 % of feed processed in DCU. Typical price of Pet coke is Rs 1300014000/MT which is far below the cost of feed processed at DCU.
Raw petcoke production at Guwahati Refinery for the last three years is as below:
Year
Production
Daily Average
(MT/yr)
production (MT/day)
Price (Rs./MT)
2011-12
61345
185.9
13829
2012-13*
55126
167.0
13956
2013-14
61151
185.3
12482
*2012-13 figures are lower since Refinery M&I shutdown was carried out.
Utilization of low cost Coal/Petcoke as fuel for power generation in place of
fuel oil
b) PETCOKE QUALITY :
The pet coke analysis report is given below:
i)
Proximate analysis :
Sl.
Tests
Units
Results
1 Moisture as received.
%wt
7.6
2 Moisture after initial drying.
%wt
0.22
3 Volatile matters.
%wt
8.8
4 Ash content.
%wt
0.28
5 Fixed carbon
%wt
90.62
6 Total Sulphur
%wt
0.68
7 Bulk Density
gm/cc
ii)
0.8584
Ultimate analysis :
Sl.
Tests
Units
Results
1 Carbon
%wt
91.19
2 Hydrogen
%wt
4.58
3 Oxygen
%wt
3.20
4 Sulfur
%wt
0.32
5 Nitrogen
%wt
0.30
6 Ash
%wt
0.21
7 Moisture
%wt
0.20
REASON FOR FINALIZING 80 TPH AS OPERATING CAPACITY OF THE BOILER :
Proposed 80 TPH capacity Petcoke boiler is due to the benefit in GRM that we can obtain by
generating maximum steam demand of the refinery from Petcoke instead of IFO since Petcoke is
cheaper than IFO.
Also, with the addition of new units like INDADEPTG, CRU and revamp of INDMAX for capacity
augmentation, power demand of the refinery is going to increase from around 14 MW to around
16.3 MW. Hence 80 TPH capacity is finalized as the capacity keeping in mind future scenario and
additional operational cushion. A summary of the work up is presented below :
Utilization of low cost Coal/Petcoke as fuel for power generation in place of
fuel oil
SUMMARY
CAPACITY, TMTPA
PETCOKE GENERATED, TMTPA
MAX BOILER CAPACITY WITH AVAILABLE PETCOKE, TPH
BOILER OPERATING CAPACITY WITH PRESENT REFINERY FUEL BALANCE , TPH
PROPOSED CAPACITY, TPH
CASE
1000
60.4
78.0
60.1
80.0
c) REFINERY FUEL MANAGEMENT POST INSTALLATION OF PETCOKE BOILER
Since the entire steam generation of GR (around 104 TPH) is presently being done from IFO (blend of
CFO + RFO), switching over to steam generation from Petcoke will lead to surplus IFO (around 18.0
TMTPA). This surplus IFO will be managed by feeding the total CFO (which is a presently blended
into IFO) to INDMAX unit to produce high value product, since INDMAX unit has undergone
capacity augmentation revamp in Feb -2016. Further, it is planned to reduce RFO generation from
DCU by implementing PDEC’s scheme of recycling the RFO back into DCU feed.
2.
BENEFIT :
a) GRM benefit possible from this project is Rs. 79.4 Crores per year, with an IRR of 30.8
%
b) Power generation cost of GR (variable cost component) is presently Rs. 10.5 per KWh.
After installation of Petcoke boiler, the power generation cost will be Rs. 6.3 per KWh
(approx) i.e, there is a reduction of Rs. 4.2 per KWh.
c) Process steam cost presently is Rs. 2348.7 / MT. After installation of Petcoke boiler, the
Process steam cost will be Rs. 1386.0 / MT (approx) i.e. there is a reduction of Rs. 962.7
/ MT
d) Presently GR TPS operates in three boiler operation philosophy (Boiler 5, 6, & 7). After
installation of Petcoke boiler, the refinery will operate in Two boiler operation philosophy
(1 Petcoke fired boiler + Boiler 6/7).
e) Petcoke price is at Rs. 12914 per MT whereas Fuel oil price is at Rs. 33792.5 per MT
(Prices are average of 2010-11, 2011-12, 2012-13). This shows that the price of Petcoke is
about 1/3rd the Price of Fuel oil. Therefore, generation of HP steam from Fuel oil is three
times more costly than generation of steam from Petcoke.
Provision of Plate type of Heat Exchanger in ARU
By – P.E.Kishore Babu, Process Manager,
M.Sankar, Chief Technical Services Manager, Mathura Refinery
Abstract:
The oil and gas industry is continuously faced with new challenges. Foremost, demand for
energy is on the increase, especially in the rapidly growing economies in India, China etc.
Secondly the development and production of oil and gas reserves are becoming more and more
difficult and risky. Also new energy deposits, such as unconventional gas or the development of
oil sands and oil shale, are becoming of greater importance. Efficient and cost-effective energy
production will be the decisive competitive factor. At the same time the top prerogative
remains safety, to ensure that personnel and environment come to no harm. This means that
all processes have to run reliably and efficiently.
Plate heat exchangers ensure low energy input, gentle product handling and an efficient
process. With plate heat exchangers up to 96 % of the energy used in the process can be
recovered.
The primary function of the Amine Regeneration Unit (ARU) is to regenerate the Rich Amine
(MDEA) received from the various secondary processing units and to supply the Lean Amine
back to these units. In ARU, Rich Amine initially received in Amine Degassing Drum and from its
bottom Rich Amine is pumped by Amine Regenerator feed pumps to Regenerator Column
through the tube side of the Amine Regenerator Feed/Bottom Exchanger. The Amine
Regenerator strips nearly all of the H2S from the Rich Amine, thus regenerating it to Lean
Amine.
The existing Feed/Bottom Exchanger i.e. Rich Amine Vs Lean Amine is designed with a heat
exchanging capacity of 7.46 G Cal/hr with approach of 30 o C. It is replaced with a welded Plate
type Heat Exchanger (PHE) for increased heat recovery from the Lean Amine which has a design
approach of 10°C. Such a lower approach is possible with PHE only as heat transfer coefficient is
very high in PHE compared to conventional Shell & Tube heat exchanger.
With the installation of PHE in place of existing Shell & Tube Heat Exchanger, the steam
consumption required at Regeneration Tower was drastically reduced by around 22-25%. At
Mathura Refinery, the savings are ~ 3500 SRFT/year and the payback period ~ 01 month. This
kind of retrofitting job of installation of PHE in ARU is 1st of its kind amongst IOCL refineries or in
Indian PSUs and will become surely a beneficial tool for reducing the operating cost of Refinery
in these days wherein the rising energy costs are denting the margins of refinery units.
Page 1 of 7
1.0 Introduction:
Plate type of Heat Exchangers (PHE) are widely used in a broad range of heating and cooling
applications in food processing, chemical reaction processes, petroleum, pulp and paper, as
well as in many water chilling applications. Some basic features of plate type heat exchangers
include high efficiency and compactness, high flexibility for desired load and pressure drop,
easy cleaning and cost competitiveness. All these qualities put together makes PHE a great
choice to be utilized in Refinery applications.
Energy is recovered in the amine recovery section, where rich amine is preheated before the
amine regenerator by means of high temperature lean amine, leaving the regenerator. More
the energy recovered in this service, lesser the energy input (inform of LP Steam) to be
provided to the regenerator re-boiler, which further reduces the operating cost of ARU.
2.0 System Description:
The primary function of the Amine Regeneration Unit (ARU) is to regenerate the Rich Amine
received from the various units viz. FCCU, DHDS, OHCU, CCRU, DHDT & PRIME-G etc. and to
supply the Lean Amine solution back to these units. The circulating Amine (MDEA) strength is ~
30 - 32 %v. In ARU, Rich Amine initially received in Amine Degassing Drum, where the amine
gets flashed & the lighters are removed. Rich Amine is then pumped by Amine Regenerator
feed pumps to Regenerator Column through the tube side of the Amine Regenerator
Feed/Bottom Exchanger. The Amine Regenerator strips nearly all of the H2S from the Rich
Amine, thus regenerating it to Lean Amine. The bottom temperature (125 ~ 130 deg C) of
Amine Regenerator is maintained using LP Steam through re-boiler.
Off-gas or acid gas generated from the top of the Amine Regenerator containing H2S is routed
to SRU-A/B/C/D. Lean Amine from the bottom of the Amine Regenerator flows to the shell side
of Feed/Bottom Exchanger for preheating the Rich Amine stream, which is ARU feed. The
cooled Lean Amine is then pumped by the Lean Amine circulation pump and flows through the
air cooled Lean Amine cooler and then through the water cooled Lean Amine Trim cooler. A
part of the lean amine further routed through Cartridge & Charcoal filters for controlling the
impurities generated in the system.
Page 2 of 7
3.0 Performance of Shell & Tube type of Feed/Bottom Exchanger in ARU of SRU
Block :
The Shell & Tube type Feed/Bottom Exchanger i.e. Rich Amine Vs Lean Amine is designed with a
heat duty of 7.46 G Cal/hr & approach of 30o C. The design and actual performance of this heat
exchanger has been as under:
Description
Lean Amine flow
Lean Amine I/L temperature
Lean Amine O/L temperature
Rich Amine flow
Rich Amine I/L temperature
Rich Amine O/L temperature
Heat Exchanged
Approach
UoM
Kg/hr x 1000
°C
°C
Kg/hr x 1000
°C
°C
G Cal/Hr
°C
Design
272
130
100
278
70
100
7.46
30
With Shell & Tube type H. Ex.
237
127
85
250
59
97
9.11
30
It may be noted that Shell & Tube type exchanger is designed with conventional approach of 30
deg C and resulting in loss of potential heat recovery.
4.0 Estimated performance of Plate type Heat Exchanger in place of Shell &
Tube Heat Exchanger:
It was proposed to install a welded Plate type Heat Exchanger (PHE) for increasing heat
recovery from the Lean Amine to have approach of 10°C against the existing 30°C. Such a lower
approach is possible with PHE only as heat transfer coefficient is very high in PHE compared to
conventional Shell & Tube heat exchanger.
The performance of the Plate type Heat Exchanger is 80% (5.98 = 13.44-7.46) more efficient
wrt heat recovery wrt design conditions & ~ 46% (4.22 = 13.33-9.11) more efficient w.r.t heat
recovery on actual conditions (Existing actual Vs Proposed actual) compared to Shell & Tube
type as under:
Page 3 of 7
Description
Lean Amine flow
Lean Amine I/L
temperature
Lean Amine O/L
temperature
UoM
Shell & Tube
based on
design flow
PHE based on
design flow
Shell & Tube
based on
actual flow
PHE
based on
actual flow
Kg/hr x
1000
272
272
237
237
°C
130
130
127
127
°C
Kg/hr x
1000
100
76
85
65.5
Rich Amine flow
278
278
250
250
Rich Amine I/L
temperature
°C
70
70
59
59
Rich Amine O/L
temperature
°C
100
119.5
97
115
Heat Exchanged
G Cal/Hr
7.46
13.44
9.11
13.33
Approach
°C
30
10.5
30
12
Additional heat
%
80
46#
recovery
# The actual heat recovery is low on account of low Rich Amine I/L temperature (Actual ~ 59°C
against the design of 70°C)
5.0 Advantages of welded Plate type Heat Exchanger against Shell & Tube Heat
Exchanger:
Welded Plate type Heat exchangers got many advantages comparable to gasketed Plate heat
exchangers as under:
S No
Shell & Tube Heat Exchanger
1.
Tubes are welded
2.
3.
4.
5.
6.
7.
8.
9.
Complex Design
Low efficiency
High Fouling, Low chocking
10-12 mm gap
Difficult to clean
High weight & Large footprints
Approach temperature less than 10-30 deg C
Maintenance – takes long time, but less
investment
Welded Plate Heat Exchanger
Fully Welded with no gaskets in between
plates
Compact
High efficiency
Low scaling , low chocking
5 - 10 mm gap
Easy to clean (all sides openable)
Low weight & Small footprints
Approach temperature 5 to 10 deg C
Maintenance – takes less
Page 4 of 7
Provision of Gasket type PHE, which is a relatively cheaper option, was also checked and same
was not considered for safety reasons, as any leak from gasket may cause spillage of Amine
along with H2S gas into the open atmosphere. Hence, the fully welded compo block type PHE
was proposed in place of Shell & Tube type Heat Exchanger.
6.0 Major hardware changes implemented in ARU:
The following changes were implemented in ARU while installing Plate type Heat Exchanger:




Provision of PHE) in parallel to Shell & Tube HX
Provision of 02 nos of PSV (01W+01S) on outlet of cold side fluid (Rich Amine)
Provision of PT & TT as per the requirement
Provision of isolations with break flanges
Page 5 of 7
7.0 Benefits:
By provision of Plate type of Heat exchanger, the feed temperature to Column has increased
from earlier 97°C to ~ 115°C. This increase in feed temperature resulted in lowering the reboiler’s Low Pressure steam (3.5 Kg/cm2g & 150 deg C). The overall savings by reduction in LP
steam is 7.0 MT/hr ~ 3500 SRFT/annum.
The typical performance of installed Plate type Heat Exchanger is as given below:
Description
UoM
Lean Amine flow
Lean Amine I/L temperature
Lean Amine O/L temperature
Rich Amine flow
Rich Amine I/L temperature
Rich Amine O/L temperature
Heat Exchanged
Approach
Additional heat recovery
Kg/hr x 1000
°C
°C
Kg/hr x 1000
°C
°C
G Cal/Hr
°C
%
Shell & Tube
based on
design flow
272
130
100
278
70
100
7.46
30
PHE
performance
-Design
272
130
76
278
70
119.5
13.44
10.5
80
PHE
performance
- Actual
277
127
67
298
61
110
13.94
17
87
8.0 Conclusion:
Implementation of Plate type of Heat Exchanger in ARU for Feed/bottom heat exchange service
is first of its kind amongst IOCL Refineries. This application has added a great value in terms of
reliability and refinery margin improvement. The investment has few months payback period.
Authors:
P E Kishore Babu is currently working as Manager (Process – Projects) and is primarily
responsible for execution of upcoming projects. He has 13 years of experience in various
Refinery Processes and Process-Projects.
Qualification: B. Tech in Chemical Engineering from Osmania University College of Technology,
Hyderabad.
M Sankar is currently working as Chief Technical Services Manager and is primarily responsible
for rendering technical services in Process Monitoring/Process-Projects. He has 20 years of rich
Page 6 of 7
experience in various fields of Refinery operations, Process monitoring, Commissioning of
Refinery Process units and Configuration studies for new refinery ,upgrading units etc.
Qualification: M. Tech in Chemical Engineering from Anna University, Chennai.
Page 7 of 7
Opportunities & Challenges for Future Purified Terephthalic Acid (PTA) Plant and
its Integration with Refinery
Kamleshwar Luckwal,Senior Production Manager, Panipat Refinery
1.0 Introduction
PTA is the preferred raw material for Polyester. PTA was first introduced in 1965 as an
alternative feedstock to Dimethyl Terephthalate (DMT) for the production of Polyester. By the
end of 1999, almost 85% of total world production was based on PTA and this continues to
grow. PTA is one of the most challenging petrochemical plants to operate due to the
following main reasons
• Use of acids: Acetic acid and HBr are used as solvent and catalyst respectively
in the oxidation process. These acids pose problems of corrosion and hazards
due to leaks.
• Slurry/solid service. The slurry and solid create lot of transfer problems and
frequent chokages are observed in the lines and equipments.
• Rotary equipments: A large number of rotary equipments are used. Many of
them are single line and hence many maintenance issues need to be addressed.
2.0 Key challenges and opportunities for future PTA plants
Elimination of HBr
Bromine compound has been added into all conventional PTA oxidation process as the
promoter to ensure the Co/Mn catalyst activity. However, the hazardous and corrosive
properties of Bromine continue to create environmental and operational problems in the
oxidation section. There is a need to work for replacement of HBr as this will eliminate the
corrosion & environmental problems and also the exotic metallurgy required due to HBr use.
This will help in reducing the investment cost of new PTA plants. Development of an alternative
simple treatment system to replace the usage of Bromine has shown progress. However, a
large scale commercial plant with such a treatment system is yet to come.
Lower Acetic acid consumption
In conventional PTA oxidation processes, the Paraxylene (PX) and Acetic Acid (HAc)
burning losses are always significant contribution to the operation cost of PTA production.
The main reason of the high PX/HAc burning loss is due to the strict 4-CBA spec (~4000ppm
max) in crude TA (CTA) after oxidation. Higher 4-CBA is undesirable because the
hydrogenation of the conventional PTA purification section is not capable of treating too high
of 4-CBA content. The severity in the oxidation reactors is reduced for higher
4-CBA in CTA. With the lower severity at lower temperature, the PX and HAc burning losses
are significantly reduced and hence significant savings in operating cost can be
achieved.
Alternative frontier is the development of a cheaper solvent to replace Acetic acid which is
being currently used in the PTA plants.
Purge stream processing
Typically, a large portion of the recovered mother liquor is recycled to the oxidation reaction
in order to recycle catalyst components to the oxidation reaction while purging a smaller
portion to a solvent recovery system so as to maintain the level of impurities and by-products
in the reaction within tolerable limits. This purge stream is incinerated in many conventional
PTA plants. A better & more economical option is to make flakes which are sourced by many
vendors for recovery of metals and benzoic acid etc. This option eliminates the incineration
of purge stream and hence environmental and ash handling issues.
There are other opportunities which need to be implemented for improving benefits. Hydrogen
recovery is one of them. H2 is used for hydrogenation reaction for purification of CTA to
improve 4-CBA in the final product. This H2 normally lends into the atmosphere. Provisions
to recover this H2 at low cost will be beneficial in future plants.
3.0 Value Addition to Refineries
PTA plants integrated with refineries provide operational flexibility & significant value addition
opportunity. With PTA plants, refineries can integrate the value chain from Naphtha to PX and
PX to PTA product. The margin in PTA varies depending on the demand and supply cycle.
However, profit of over Rs. 5 per kg of PTA is quite normal.
PX unit for producing PX feed for PTA, normally consists of Continuous Catalyst
Regeneration (CCR) unit which produces reformate. This reformate provides a very good
option for gasoline (MS) blending also. Refineries integrated with PTA unit can leverage the
advantage of this flexibility between MS and PTA mode based on the economics of PTA over
MS.
4.0 Further Opportunities for Refinery
In addition to increased product value addition, the major opportunity in new PTA plant lies in
steam Integration with the refinery.
Normally, the potential to generate low pressure or extra low pressure steam (< 2.0 kg/cm2g) in
refinery remains unutilized due to lack of consumer for the same. However, in future PTA unit
integrated with refinery, generation of low pressure/extra low pressure steam from waste/low
temperature heat recovery will help in huge energy savings as this low pressure steam
generated in refinery, can easily be utilized in steam turbine driving the process air compressor
(PAC) used in PTA to feed air to the Oxidation reactors. PAC consumes large power (typically
for a 0.5 MMTPA unit, it can consume power up to 29 MW). This machine is normally driven by
steam turbine and the turbine can be designed to consume large quantity of low pressure steam
even up to 1.0 kg/cm2g. This provides significant opportunity to refineries considering new PTA
unit to recover low temperature heat by generating low pressure steam which can be fed to the
compressor’s steam turbine drive.
Other opportunities include use of waste N2 (rich in Oxygen) from N2 unit for Oxygen
enrichment of process air to be used in PTA oxidation unit. The off gases from Oxidation section
of PTA contains high N2 (approximately 94% Vol) which can also be an area of opportunity for
recovery of N2/CO2.
1
Optimizing operating parameters by predicting feed quality
Authors- Sh Mousom Some (STSM-PNC)
Sh Puranjay Choudhury (AMPS-PNC)
Abstract
Purpose- The purpose of this paper is to study the methodology for optimizing unit
operations by predicting feed quality.
Synopsis- The economics of operation in a Naphtha Cracker Plant depends a lot on the nature
of the feed Naphtha which is used for cracking. Naphtha is sourced from refineries for
processing at Naphtha Cracker and vary widely wrt quality.
In the absence of a Real Time Optimizer, it becomes difficult for the operating personnel to
assess the Naphtha quality on real time and adjust the operating conditions for optimum
Olefin yields. The actions taken under the above circumstances are mostly post facto in
nature, by which time we have already lost some yields on olefins.
Prediction of Naphtha quality through calculation of the expected time of feeding followed
by creation of Paper blends can, to a substantial extent, assist the operating personnel to tune
the operating conditions based on the predicted Naphtha quality, so as to have the most
optimum and economic product slate.
The expected Naphtha quality at different time periods is run on a virtual environment i.e in
PYPS software to estimate the product yields based on varying Naphtha quality. By varying
the severity of operation at different time periods, the most optimum product mix can be
obtained.
The offline simulation in PYPS software also gives an idea on the effect of severity on the
run length of the coils. This shall further help the operations personnel to select the most
optimum range of severity so as to have the right balance between maximizing Olefin yields
and the coil run length.
Introduction and Objectives
In the highly competitive petrochemical market, the key to have an edge over the competitors
rests upon innovative and price effective ways of operation. With rising feed stock costs and
shrinking margins, the thrust should be on optimizing the unit operations in such a way, so as
to reap the maximum benefits out of them.
This basic idea had triggered the thought process for the paper. With proper implementation
of the model which will be detailed out in the subsequent sections of the paper, a typical
Naphtha Cracker can get the most optimized product slate, along with enhanced period of
operations with minimum downtime.
2
Optimizing operating parameters by predicting feed quality
A typical Naphtha Cracker of an integrated Refinery & Petrochemical Complex faces the
problem of varying Naphtha qualities. Naphtha is pooled from multiple refineries having
different configurations and Crude mix sourcing.
This results in wide variation in the quality of the Naphtha which is fed to Naphtha Cracker.
The Naphtha qualities vary mainly wrt PIONA and Distillation. The biggest challenge before
a Panel Engineer then becomes, to adjust the operating conditions very frequently to achieve
the target yields of olefins. The actions taken mainly w.r.t. the severities or Coil Outlet
Temperature are post facto in nature, which may not be the most optimized. Additionally
unidirectional adjustment of severity to achieve target yields may not always be economical
considering the long term impacts on the heater coil run lengths.
The tool developed for prediction of Naphtha feed quality and adjustment of operating
parameters based on the predicted Naphtha quality at different timelines, shall not only help
in maximizing olefin yields but will also address the adverse impact of deteriorating heater
coil run lengths, thus striking the right chords between economics and equipment longevity.
Related Work
A lot of work has been done to study the impact of various Naphtha feeds on Olefin yields. In
this case PYPS (Pyrolysis Yield Prediction Software) has been used to predict the product
yields on Naphtha of different qualities.
Since Naphtha of different qualities are pooled in a single tank, the quality of the blend
especially the distillation data have been established in UNISIM Simulation Software.
System Model
The system model requires the qualities of the various Naphtha batches and the tank where
they are planned to be received. In case of 2 different Naphtha batches, a blend is generated
based on the respective batch sizes, in UNISIM software (Assays are fed in “Oil Manager” of
UNISIM software, providing the details on distillation data. Multiple assays are made to form
a blend). The expected times for feeding the blends into the heater are calculated based on the
line size and the distance till the feed inlet point.
Based on the feed qualities and the tentative feeding timelines, the olefin yields are calculated
using PYPS at different severities. Apart from the yields, PYPS also calculates the tentative
heater run length for a particular feed.
The cracking margins against the various severities in the different time periods are
calculated thus providing a comparative information on the most economic mode of operation
in any particular time periods.
Panel engineers are accordingly advised to operate at the most optimized severity which
merges both product yields as well as heater run lengths.
3
Optimizing operating parameters by predicting feed quality
Conclusion
This tool shall help in adopting a strategic rationale behind operating cracker at different
severities which shall be predictive in nature and backed by economic implications.
ABSTRACT
SWOT analysis and metamorphic
change towards Energy &
Petrochemicals from standard fuel
complex
FUTURE INDIAN
REFINERY
CONFIGURATION
Future Indian Refinery Configuration
– SWOT analysis and metamorphic change towards Energy & Petrochemicals from
standard fuel complex
Srinivas Moturi, Deputy Manager, MRPL, India
Nagaphani Kumar Ravuri, Deputy Manager, MRPL, India
Change is always an opportunity and a threat. Glo alizatio , u til last de ade the Left’s a ti hrist, is
now the same thing but for the Right. Polarization of political ecosystems in Western nations has
produced and will continue to have a huge impact on the globalization process, world politics and
economics. Even though the world had been forewarned by the referendum in U.K., BREXIT still
brought about unprecedented consequences. A successful Brexit will no doubt encourage such
sentiments to grow. In sum, Europe and the U.S. are experiencing a fundamental shift in social
physiological framework and political ecosystems. From a strategic and global perspective, Brexit may
be defined as the first wave of anti-globalization and rising populism that washes over the world, the
advanced nations.
In these days of political turmoil and de-globalisation, a ou tr ’s usto ised sustai a le usi ess
models play a fundamental role in the Nations stride towards global excellence. An integrated fuel,
energy & chemical policy of the country is second to none in the present-day nation’s growth. Human
development index is directly correlated to availability of energy and chemical, flooring a path forward
for the nation
India, the second largest populated nation in the world with more than a billion people has an economy
which is growing at nearly 8% over the last decade and about 6% on the average since her
independence in 1947. It is expected that India's economy will go at the same rate even till 2050, which
will naturally demand enormous amounts of energy and supply of basic chemicals for indigenous
growth.
Though India is presently the fourth largest electricity producing country in the world, her per capita energy
consumption (500 kWh) is rather small, which is only about 1/2 of China, ¼ th of World average and about
1/13th of developed nations. However, India aspires to reach at least the global average by 2050, which
would require her to produce about 1300 GW of electricity, ten times more than the present value of about
130 GW. Next to energy, Chemicals and petrochemicals are area of concern for the nation. We have a huge
supply and gap in polymers, fibre inter ediates, s theti fi re a d elasto ers for proje ted atio s’
growth. Per capita plastic consumption in India is still hovering at 7.0 kgs as compared to 46 kgs in China
and 65 Kgs in Europe. This signifies huge potential for future growth going by current global average per
capital consumption. Petrochemicals sector is one of the fastest growing segments with a growth rate of
13 per cent which is more than twice of growth of India's gross domestic product (GDP) and the global
growth rate in petrochemical space which is stagnant at 6 per cent. The outlook is also stable and the
chemicals market is expected to grow at 11-13% p.a. over the next decade
One the other side for Indian refiners, Global surplus refining capacity is expected to rise by more 1 million
b/d to 5.3 million b/d in 2021 creating "significant pressure" on refining margins in the medium term. The
IEA estimates that almost one-sixth of global oil demand will be met by fuels by-passing the refining sector
such as biofuels and NGL's in 2021, which will exacerbate the surplus capacity issue. We as a nation became
net exporters of diesel and petrol. Any new requirement for fuel growth will be catered by capacity
augmentations and de bottlenecking refinery process units. Also, India ratified the Paris agreement COP-
21 o li ate ha ge. I dia, hi h is the orld’s fourth-largest carbon emitter accounting for 4.1% of the
total global emission, is the 62nd nation to ratify the agreement. Bio-fuels and renewables are a challenge
for the refiners.
Electric cars could upend oil markets much sooner than everyone thinks., the rapid decline in the cost of
building batteries for electric vehicles (EVs) will make them cheaper than the internal combustion engine
in just a few years. By the 2020s, EVs could beat conventional vehicles on price, a shocking development
and a potential epochal shift for energy markets. Battery prices declined by 35 percent in 2015, another
impressive feat for the technology as it marches towards both relevance and market share. EVs – on an
unsubsidized basis – will be just as affordable as a car that runs on gasoline within few years. That means
that
, EVs ill rea h the poi t of lift-off for sales. The ost-competitive prediction for EVs even
assumes that gasoline-powered cars continue to improve efficiency at a rate of 3.5 percent per year.
EVs will control 35 percent of the auto market by 2040. Behind those cost declines is the dramatic fall in
the prices of batteries. Costs for lithium-ion batteries have plummeted to just $350 per kWh in 2015, a 65
percent cost reduction since 2010. But battery manufacturers are not done yet. By 2030, costs will fall to
$
per kWh, or less tha half of toda ’s le els. Fro there, osts ill o ti ue to de li e.
For Refiners, the conclusions should be alarming. The electric vehicle revolution could turn out to be more
dra ati tha go er e ts a d oil o pa ies ha e et realized. - Bloomberg news
Projected that EVs displacing 2 million barrels per day (mb/d) of oil demand as early as 2023. That is just
the start. The real pain will come after that point as EV sales start to skyrocket. EVs could capture 35 percent
of the market by 2040, which would displace 13 mb/d. For an oil market, currently in tatters because supply
is exceeding demand. EVs present an existential threat.
On date, Indian peninsula has some constrained factors like Saturated refining capacity, COP-21 Ratified
Paris agreement on emissions, volatile Crude prices, de globalisation geo political factors, evolving EV
market. But low per capita energy and low per capita polymer usage paves a path forward for Indian
refiners. In contrast to the world, Nation can cherish the advantages of domestic factors and consolidating
Make in India theme is reviewed with SWOT analysis. The strategic refinery configuration is derived from
metamorphic change towards Energy & Petrochemicals from standard fuel complex. The configuration
includes Petrochemicals, sustainable technologies, cleaner fuels and the theme of Make in India and
Make for India
Precursors for future Indian refiners
A. Petrochemicals - Economical vector for the configuration
Methanol:
The methanol market in India is projected to grow at a CAGR of over 7% till 2025. Growing demand
for methanol as a raw material for manufacturing of formaldehyde coupled with its applications in
the manufacturing of acetic acid, methyl amines, MTBE, etc. are some of the major factors
expected to aid the India methanol market during forecast period. Increasing use of methanol in
the ou tr ’s phar a euti al se tor alo g ith ide tifi atio of o el e d uses of etha ol are
contributing to the demand for methanol in India. Growing applicability of methanol to
manufacture acetic acid is also expected to aid the market of methanol in the country during 2016
- 2025. Moreover, majority of the demand for methanol in India is being catered by imports
MEG:
Mono ethylene glycol, commonly referred to as ethylene glycol, or MEG, is the largest produced
glycol by volume, amongst all glycols. Almost all the consuming sectors of MEG have witnessed
growth rates of more than 10% in the past 5 years, indicating a strong demand of MEG in the
present market scenario. Polyethylene Terephthalate or PET, is the largest consumer of MEG,
accounting for more than half the domestic production of the country and expected to achieve
growth rates close to 20% in the next 5 years. The textile and pesticide industries are other major
consumers of MEG in the country and they are expected to witness modest growth in the coming
years. The upsurge in the polyester fabric industry mandated an incessant supply of MEG which
further fuelled its demand.
Though production has seen a modest growth, it is not enough to meet domestic demand. India's
limited domestic production mandates users and distributors to procure MEG through imports.
Imports are at their highest during peak summer months due to its primary application in the
manufacture of PET which is used in the packaging of soft drinks and other beverages that show
maximum sale in summer.
Ethylene:
Polyethylene, also called polythene, is the worlds most widely used plastic, primarily used to make
fil s used i pa kagi g a d plasti ags. Pol eth le e o su es ore tha half of the orlds’
supply of ethylene, derived from various petrochemical olefins. India leading Asia Pacific is
expected to see the largest plastic polyethylene (PE) demand growth from 3.6 million tons (mt) to
8.2 mt in the next 10 years. According to , Platts Polyethylene Report – Glo al Outlook to
,
I dia’s PE de a d is likel to i rease
%, far surpassi g Asia’s proje ted gro th rate of %
a d Chi a’s % for the period.
Demand for high density polyethylene (HDPE), used in the manufacturing of such things as plastic
sheeting for ducting and appliance hoods, is expected to climb to more than 4 million mt by 2023,
while demand for linear low density polyethylene (LLDPE), used in the making of such things as
industrial containers, is projected to approach 3 million mt over the period. Demand for low density
polyethylene (LDPE), used in the making of such things as kitchen cutting boards, is expected to
reach 1.2 million mt. If proje tio s are orre t, I dia ill e the se o d largest i porter of
pol eth le e
, ehi d Chi a, I dia’s e pe ted pol eth le e deficit will be 3.4 million mt
, ehi d Chi a’s likel
illio
t shortfall a d ahead of defi its fore asted for Afri a a d
Europe. A d hile Chi a’s de a d is still learl outpa i g I dia’s, the latter is pro i g to e a
influential customer in the growing polyethylene markets, during a decade outlook period where
global oversupply is becoming more of a concern, he said.
Acetic acid:
The acetic acid market in India is anticipated to grow at a CAGR of over 9% during 2016 - 2025,
because growing utilization of acetic acid as a raw material to produce Purified Terephthalic Acid
(PTA) coupled with increasing applications in manufacturing of ethyl acetate, butyl acetate and
anhydrides in the country. Acetic acid is used in a wide range of end user industries like
pharmaceuticals, textiles, packaging, films, adhesives, paints & coatings, etc. Acetic Acid is one of
the major organic chemicals imported in India. Almost all the consuming sectors of Acetic Acid have
witnessed growth rates of more than 10% in the past 5 years, indicating a strong demand of Acetic
Acid in the present market scenario. Purified Terephthalic Acid (PTA) is the largest consumer of
Acetic Acid, accounting for more than half the domestic production of the country and expected to
achieve growth rates close to 20% in the next 5 years. Esters, Camphors and Dyes are other major
consumers of Acetic Acid in the country expected to witness modest growth in the coming years.
Western region of the country is the largest hub of Acetic Acid consumption, accounting for more
than half of domestic Acetic Acid consumption. Other regions of the country share approximate
similar Acetic Acid requirements, which are however, expected to increase in the future.
In 2015-2016, production of Purified Terephthalic Acid (PTA) in India stood at 6,950 Kilo Tons, and
this is expected to increase steadily during the forecast period. As a significant amount of acetic
acid is consumed during the manufacturing of PTA, expanding PTA capacity is expected to positively
influence the India acetic acid during 2016 - 2025. In addition, PTA is also used during the
manufacturing of polyester fiber in India. In 2015, the production of man-made fiber in India stood
at 4,140 thousand tons. Ethyl acetate is forecast to be fastest growing application segment
in India acetic acid market during the forecast period.
Maleic anhydride:
The maleic anhydride market size, in terms of value, is projected to reach USD 5.08 Billion by 2020,
at a CAGR of 6.8% between 2015 and 2020. The market is driven by the growing automotive
industry and rapid industrialization in India. The increasing demand of automobiles in emerging
economies such as India is the driver for the product. Another major factor contributing to the
growth of maleic anhydride is rising global demand for UPR, and 1,4-BDO
Styrene:
Styrene monomer is an important feedstock for the manufacture of several industrially valuable
chemicals such as ABS (acrylonitrile butadiene styrene), SAN (Styrene acrylonitrile), styrene
butadiene and in the production of polystyrene.
The latter is the most important application of styrene in the country. Most of the produced styrene
monomer is used in the manufacture of mundane items such as CD cases, disposable cups, food
containers and refrigerator door liners. The northern region leads domestic demand for styrene in
the country, followed by the western and southern regions.
The demand for styrene has significantly increased since 2010, after the chemical market recovery
from the global economic slowdown. The entire demand for styrene is met through imports due to
lack of any production facilities within the country. This makes the demand susceptible to
international fluctuations in prices as well availability of styrene. Due to lack of production facilities,
India is entirely dependent on production capacities in other countries to meet its growing styrene
demand. An abrupt change in international prices therefore, harshly affects domestic users of
styrene in the country. Since there are no plans of any capacity initiations in the near future,
international producers hold an extraordinary grip on the domestic market.
Styrene is relatively inexpensive to move and hence, widely exchanged between different regions.
Asia currently accounts for more than half of global styrene demand and is expected to remain the
global styrene growth driver. Continued industrial development, population growth and rising
income levels are key drivers, all of which are dominant in the upcoming Asian markets and in India.
In the global market, styrene demand remains dominated by its main derivative, polystyrene, and
this industry has reached maturity in several developed countries of the west. Styrene demand in
the Americas and Europe has, therefore, declined in recent years due to reduced levels of
manufacturing and material substitution in packaging. Other important applications of styrene in
the global market include production of acrylonitrile-butadiene-styrene (ABS) and styreneacrylonitrile (SAN) resins, styrene-butadiene (SB) copolymer latexes and unsaturated polyester.
A ordi g to I dia Ther oplasti Elasto ers Market B T pe, B E d Use Appli atio ,
Competition Forecast & Opportunities, 2011, the arket of ther oplasti elasto ers i
India is anticipated to grow at a CAGR of around 7% during 2016 – 2025. Styrenic Block Copolymers
dominated India thermoplastic elastomers market in 2015, and the segment is expected to
continue dominating the market through 2025 owing to its wide range of applications in various
industries
PVC:
I dia PVC Market Stud ,
, the PVC arket i I dia is proje ted to e hi it a CAGR of
over 14% during 2016 – 2025, on account of expanding agriculture sector, coupled with strong
growth in infrastructure development, rapid industrialization and rising urbanization. Moreover,
increasing favourable government policies and investments, growing population and rising
construction activities in the country, are further expected to drive consumption of PVC in India
during the forecast period.
PVC is tailored for utilization as a common plastic material in the construction sector,and is used in
window frames and shutters, pipe cabling and coating, etc. PVC is a preferable plastic material for
broad array applications including flooring, wires & cables, films & sheets, etc., on account of its
high rigidity and flexibility. Moreover, PVC also offers fire resistance owing to its high chlorine
o te t, hi h is further oosti g PVC’s o su ptio i the o stru tio se tor. As per 12th fiveyear plan (2012-2017), a total investment of USD1,000 billion is planned to be made in the
infrastructure sector of the country, which is further expected to drive India PVC market during the
forecast period.
List of major petrochemicals and their market study
S. No
Product
Existing
Domestic
Capacity (KTA)
Demand
(KTA)#
Net imports in
2015-16 (KTA)
Value
%CAGR
(Rs Crore)
(2015-25)
1
Methanol
492
1900
1668
2771
5.7
2
PVC
1482
2775
1498
8756
6.5
3
PTA
6570
4800
524
2508
8.3
4
MEG
1215
2350
1040
4876
9.5
5
HDPE
1670
1968
805
7012
6.7
6
Acetic Acid
433
847
785
2011
4.2
7
Styrene
0
717
717
5158
5.2
8
LLDPE
1670
1430
565
4378
6.7
S. No
Product
Existing
Domestic
Capacity (KTA)
Demand
(KTA)#
Net imports in
2015-16 (KTA)
Value
%CAGR
(Rs Crore)
(2015-25)
9
EDC
190
720
584
1019
6.5
10
VCM
906
1100
349
1672
6.5
11
LDPE
205
535
351
2870
6.1
12
Toluene
258
468
339
1605
7.0
13
SBR
270
283
138
1286
7.5
14
Phenol
80
272
242
1428
8.6
15
Polyols
155
274
165
2162
9
16
Vinyl
0
Acetate
Monomer
(VAM)
Ethyl Vinyl 15
Acetate
(EVA)
140
140
869
5.8
174
164
1605
8.9
18
LAB
530
590
206
1675
3.7
19
PP
4640
3840
17
6.7
*Net imports of petrochemicals has increased from Rs.10,601 Crore in 2007-08 to Rs. 53,661 Crore in
2015-16
*Source: Department of Commerce, Ministry of Commerce & Industry #2015 estimated demand
B. Bio fuels - Sustainable vector for the configuration
The global biofuels market as alued at $
.
illio €
.
i
a d is proje ted to
reach $246.52 billion by 2024 at a CAGR of 4.92%. Market Resear h’s e stud predicts biofuels
and bioenergy are a critical part of the renewable energy mix, being the only renewable energy
that can be used across all three energy sectors (electricity, heat, and transportation). The global
bioethanol market is experiencing tremendous growth due to its increased use in various end user
industries, specifically in the transportation sector, and its usage is being encouraged due to the
ban on 100% petrol in many countries. It is replacing other conventional fuels and being used as an
additive along with petroleum products.
The volatility in crude oil prices and increased greenhouse gas (GHG) emissions are encouraging
government institutions to promote consumption of bioethanol. The sources of Bio fuel to a
growing economy like India are
•
•
•
Algae oil – the ethanol source from refinery
Hydrothermal liquefaction
2G ethanol
Algae oil – the sustainable Ethanol source from refinery
CO2 enriching techniques from refinery Flue gas steams of Fluidised Catalytic Cracking (FCC),
Captive Power Plants (CPP) flue gases and Process flow diagram for stage wise implementation in
existing refineries, Algae culture to produce diesel, Aviation fuel, Gasoline and Ethanol for the
refiners are elaborated.
SUNLIGHT
FCC FLUE GAS
CO2
ENRICHING
CPP FLUE GAS
SEAWATER
ALGAE
CULTURE
BIOFUELS
ATF
GASOLINE
DIESEL
ETHANOL
Use of Algae Ethanol for firing the process heaters with or without blending with HFO is being
reviewed. This can lead to lower SOx and HC emissions, while the flue gas generated can be further
sent downstream to Algae Ethanol system as a source of CO2. The surplus refinery fuel gas can be
sent for production of value chain petrochemicals.
C. Aviation fuel – On demand product for future
I dia’s i il a iatio i dustr is o a high-growth trajectory. India aims to become the third-largest
aviation market by 2020 and the largest by 2030. The Civil Aviation industry has ushered in a era of
expansion, driven by factors such as low-cost carriers (LCCs), modern airports, Foreign Direct
Investment (FDI) in domestic airlines, advanced information technology (IT) interventions and
growing emphasis on regional connectivity. India is the ninth-largest civil aviation market in the
world, with a market size of around US$ 16 billion. India is expected to become the third largest
aviation market by 2020.
Indian domestic air traffic is expected to cross 100 million passengers by FY2017, compared to 81
million passengers in 2015, as per Centre for Asia Pacific Aviation (CAPA). India is among the five
fastest-growing aviation markets globally with 275 million new passengers. The airlines operating
in India are projected to record a collective operating profit of Rs 8,100 crore (US$ 1.29 billion) in
fiscal year 2016
D. Crack ratio – most critical for future
Future Indian refiners are to be designed to crack Hydrocarbon chains as per market demand. The
majority of Straight run streams are to be diverted to intense cracking units like PFCC, Steam
cracker to produce in demand C2, C3, C4 olefins. The future Indian refiners will be energy intense
complex configurations with highly cracked output.
STRENGTHS
WEAKNESS
NAPHTHA AVAILABILITY
IMPORTED TECHNOLOGY
SKILLED MANPOWER
IMPORTED MACHINERY
AVIATION FUEL CONTRACTS
LACK OF IPR, R&D - CATALYSTS
GEOGRAPHICAL PRESENCE
i
OPPORTUNITIES
THREATS
GROWING GDP
ELECTRIC VEHICLE
LOW PER CAPITA ENERGY
COP-21
LOW PER CAPITA POLYMER
RENEWABLE STATUTORY OBLIGATION
INFRASTRUCTURE PLASTICS RAISE
GEOPOLITICAL DEVELOPMENTS
AVIATION FUEL DEMAND
Future Indian refinery schematic- Next generation from present fuel complex
METHANOL
ACETIC ACID
C1
ETHYL VINYL ACETATE
FUEL GAS
VINYL ACETATE MONOMER
PVC
PE
MEG
C2
VCM
EDC
LPG
MA
C4
C3
STEAM
NAPHTHA
CRACKER
AVIATION FUEL
KERO
LAB
C6
GASOIL
C2 =
VGO
DOWN FLOW
PETRO FCC
EB
PP
C3 =
POLYOLS
PHENOLS
NAPHTHA
AROMATIC
COMPLEX
NAPHTHA
VGO
GASOIL
CDU
SLURRY
HYDROCRACKER
PX
21st Refinery Technology Meet (RTM), 2017
Fluidized bed gasification: Technology status, Challenges and Modeling approaches
Ankit A. Jain and Ajay Gupta
Refining R&D, Reliance Industries Limited, Jamnagar-361142
Introduction: Coal and Biomass Reserves in India
The energy situation in India is a paradox wherein though it has abundant reserves of coal and
biomass, yet it is energy deficit. India boasts of 7.1 % of world’s total coal reserve (British
Petroleum 2010), 70 % of the total power generated in India is from coal-fired power plants (Reddy
et al. 2013). Various organizations (Central Electricity Authority 2014; Central Electricity
Regulatory Commision 2011) have reported that India faces 10-13 % deficit in terms of energy
supply to demand. With its relatively comfortable resource base compared to limited known oil &
gas resources; coal is the obvious, affordable and sustainable choice for generation of electricity.
Therefore, India’s power development programme is heavily dependent on coal and its quality is
an important parameter that influences the performance of the power stations. Conventional
technologies based on the coal combustion process are low on efficiency and release greenhouse
gases such as carbon dioxide, sulphur dioxide and nitrogen oxides. Similarly with respect to
biomass, reports have shown that India produces 686 MT of biomass and 34% of the same is
surplus and can be used to produce bio-energy. It is estimated that 17% of India’s total primary
energy can be fulfilled through energy produced with biomass (MNRE 2017). India has a separate
ministry i.e. Ministry of New and Renewable Energy (MNRE) wherein it provides various forms
of subsidies and incentives to biomass power projects. The number of projects wherein Biomass
power is being generated has been listed in Table 1.
Looking at the scenario where high demand of energy is forecasted, it is essential that India looks
toward other technologies also termed as clean technologies. Development of means to convert
coal/biomass from its native form into useful gases and liquids in ways that are energy efficient,
non-polluting and economical is key in fulfilling the needs of our society. Gasification has been
internationally regarded as an effective way for clean use of coal/biomass especially for the
production of energy and production of synthetic chemicals.
1
21st Refinery Technology Meet (RTM), 2017
Table 1: State-Wise/Year-Wise List Of Commissined Biomass Power/Cogeneration Projects (As
On 01.04.2016) (IN MW) (MNRE 2017)
S.No.
State
Upto
31.03.2012
2012-13
363.25
17.5
380.75
15.5
27.92
43.42
2
Andhra
Pradesh
Bihar
3
Chattisgarh
249.9
4
5
6
Gujarat
Haryana
Karnataka
Madhya
Pradesh
20.5
35.8
441.18
8
1
2013-14
2014-15
2015-16
Total
15
15
279.9
10
9.5
50
13.4
12.4
112
111
56.3
45.3
872.18
8.5
7.5
10
9
Maharashtra
603.7
151.2
185.5
184
9
10
11
Odisha
Punjab
Rajasthan
20
90.5
83.3
34
10
16
8
15
7
12
Tamil Nadu
532.7
6
32.6
31.6
39
626.9
13
Uttarakhand
10
20
20
13
50
14
Uttar Pradesh
644.5
132
93.5
842
15
West Bengal
16
10
Total
3135.33
465.6
7
158
35
96.38
1220.78
20
155.5
108.3
26
412.5
405
400
4831.33
Research and development on gasification with high ash coal (Indian coal) and biomass needs to
be a priority area. There are various technologies developed and used worldwide for coal
gasification including moving bed , fluidized bed (bubbling and circulating fluidized bed) and
entrained bed gasifier being the prominent ones (Gräbner 2015). Moving bed and the fluidized bed
are considered more apt for handling high ash coal (Collot 2006) among others. Fluidized bed has
certain advantages over moving bed in terms of scaling up and environmental issues. Moving bed
gasifier generate tarry products whereas fluidized bed gasifier yield only gaseous product as the
volatiles get cracked up facilitating more environment friendly products and also easier plant
operation (Kristiansen 1996). The advantages of fluidized bed gasifier are well documented (André
et al. 2001). Good gas solid contact, excellent heat transfer characteristics, better temperature
control, large heat storage capacity, good degree of turbulence and high volumetric capacity are
2
21st Refinery Technology Meet (RTM), 2017
few prominent advantages of the fluidized bed. There are various sub-groups into which the
fluidized bed technology systems can be fit into essentially atmospheric/pressurized systems,
bubbling/circulating fluidized bed. Each of these type of fluidized bed have their own set of
advantages in terms of: scale, operability, co-gasification, economics and yields.
Product Gases Out
Bubbling fluidized bed (BFB) is considered apt for
gasification of high ash coal and biomass (Collot
Particle Level
Cyclone
Solid-Gas Reactions
Pyrolysis, Combustion,
Gasification
Freeboard
2006) (see Figure 1). Understanding of the gas-solid
hydrodynamics, chemical reactions and the effect of
Particles/Fly Ash
various operating parameters on the performance of
the gasifier plays a key role in the design,
Bubbles
optimization and scale up of the plant. Mathematical
models
are
pivotal
in
developing
Fluidized Bed
these
Solids Feed
understandings at a much reduced temporal and
Bottom Ash
financial efforts in comparison to the conventional
Inlet Feed Gas
(Air/O2/Steam)
approach of conducting comprehensive experimental
investigations from laboratory scale test unit to a pilot scale
Figure 1: A schematic representation of a bubbling
fluidized bed gasifier
plant, before building a full scale commercial demonstration
plant (Hamel & Krumm 2001).
Approach to model the reactor
There are three main approaches to model a fluidzed bed reactor, each of the approach has its own
set of advantages and disadvantages. A summary of the essential aspects of these approaches is
shown in Table 2. We applied all three approaches to model the case of a fluidized bed coal
gasifier. The exercise was useful in developing a strategy to model the reactor. The three main
main approaches are:

Data driven models: Artificial neural network, Regression based model, etc.

Equilibrium based models: Stoichiometric and non-stoichiometric

Rate based models: Computational Fluid Dynamics, Kinetic model
There are various other options to model the reactor which may be combination of the above
mentioned approach i.e. reactor network models, equilibrium and kinetic model, etc.
3
21st Refinery Technology Meet (RTM), 2017
Table 2: Approaches to model a reactor (case study: BFBG, adapted (Gómez-Barea & Leckner 2010))
Approach
Data Driven
Models
Essential aspects
No governing equations
Results
Outlet gas
composition,
temperature and
carbon conversion
Advantages
No understanding of
any complex process
in the reactor is
required
Thermodynamic
models
1. Based on Gibb’s
minimization theory
2. Calculates equilibrium
gas composition and
temperature
1. Based on first
principles, solves the mass
and energy balance
equations
2. Instead of solving the
momentum equations,
semi-empirical
correlations are used
3.Ideal reactors/
combination of ideal
reactor assumptions for
phases (CSTR, PFR,
compartments)
1. Mass, energy and
momentum balance
equations are solved
2. Constitutive relations
and closure laws are
adopted
Outlet gas
composition,
temperature and
carbon conversion
No understanding of
any complex process
in the reactor is
required
Profile of gas/solid
species composition,
temperature and
hydrodynamics
across the reactor
1. Computationally
less intensive than
CFD models
2. Can be used to
study the influence of
various input
parameters on the
performance of the
reactor
3. Gives sufficient
details for engineering
applications
Profile of gas/solid
species composition,
temperature and
hydrodynamics
across the reactor
Useful for exploring in 1. Computationally very
depth hardware details expensive and time consuming
solution
2. Uncertainty of various
parameters
3. Availability of in-depth
experimental data to validate the
model results
Chemical reaction
engineering
models (CRE)
Computational
Fluid dynamic
(CFD) models
4
Disadvantages
1. No insights into the complex
relationships between the input
and the output parameters
2. Quality and quantity of data is
required for building a robust
model
No insight into the
hydrodynamics of the reactor
Types
1. Regression analysis
2. Artificial Neural
Networks (ANN)
3. Fuzzy Rule Based
Systems (FRBS),etc.
Flow structure and the range of
applicability depends on the
correlations used
1. Davidson Harrison
Model (DHM)
2. Kunii Levenspiel
Model (KLM), etc.
1. Equilibrium Models
(EM)
2.Modified Equilibrium
Models (MEM), etc.
1. Eulerian Eulerian
model (EEM)
2. Eulerian Lagrangian
model (ELM), etc.
21st Refinery Technology Meet (RTM), 2017
Summary and Conclusions:
1. Biomass and coal reserves in India are crucial to meet its energy demand
2. Fluidized bed gasification technologies is a promising route to produce clean fuel
3. Government of India is giving subsidies and incentives to projects wherein power is to be generated
through biomass energy
4. Catalytic fluidized bed gasification are important routes to lower the tar and increase the efficiency
of the process
5. Research in the area of fluidized bed gasification especially at pilot/demo scale is required to
generate data and develop the understanding of the process
6. ANN is useful when a large amount of quality data is available to train and test the neural network.
It is also useful when sufficient details of the hydrodynamics and the kinetics involved in the reactor
are not known. It is not useful for scale-up studies of a reactor.
7. For reactors where high residence time and rapid chemical reactions are seen, an equilibrium model
may be sufficient to model the reactor. In the case of BFBG, an equilibrium model did not show a
good match with experimental data, but it did provide a fair indication on the operational limits. It
is also useful in understanding the qualitative change with alterations in input parameters.
8. Results with rate-based models, where the hydrodynamics and the chemical reactions in the reactor
are taken into account, best described the experimental data of BFBG. These models provide
sufficient knowledge for the scale-up and optimization of the reactor.
9. CFD models are computationally intensive, especially for the simulation of commercial-scale
reactors. However, with the availability of cheap computational power, use of this tool to simulate
reactors is increasing. Models where flow and mixing knowledge obtained from a CFD model are
utilized in quantifying flow and mixing in a CRE model have been reported in literature.
5
21st Refinery Technology Meet (RTM), 2017
References:
André, F., Fernandes, N., & Lona, L. M. F. (2001). Fluidized-bed reactor modeling for
polyethylene production. Journal of Applied Polymer Science, 81, 321–332.
British Petroleum. (2010). BP Statistical Review of World Energy. Retrieved December 19, 2016,
from www.bp.com/productlanding.do?categoryId=6929&contentId=7044622S
Central Electricity Authority. (2014). Load Generation Balance Report 2014-15. Retrieved
December 29, 2015, from http://www.cea.nic.in/reports/yearly/lgbr_report.pdf
Central Electricity Regulatory Commision. (2011). Annual Report. Retrieved February 04, 2016,
from http://www.cercind.gov.in/annual_report.html
Collot, A.-G. (2006). Matching gasification technologies to coal properties. International Journal
of Coal Geology, 65(3-4), 191–212.
Gómez-Barea, A., & Leckner, B. (2010). Modeling of biomass gasification in fluidized bed.
Progress in Energy and Combustion Science, 36(4), 444–509. doi:10.1016/j.pecs.2009.12.002
Gräbner, M. (2015). Industrial Coal Gasification Technologies Covering Baseline and High-Ash
Coal. Weinhein, Germany: Wiley-VCH Verlag GmbH & Co.
Hamel, S., & Krumm, W. (2001). Mathematical modelling and simulation of bubbling fluidised
bed gasifiers. Powder Technology, 120, 105–112.
Kristiansen, A. (1996). Understanding of coal gasification. London: IEA coal research.
MNRE.
(2017).
Retrieved
April
4,
2017,
connected/biomass-powercogen/
6
from
http://mnre.gov.in/schemes/grid-
Paper for RTM 20–22 April 2017, Visakhapatnam, India
FIRST PRINCIPLE APPROACH TO DEBOTTLENECK PROPYLENE
RECOVERY UNIT TO PRODUCE POLYMER GRADE PROPYLENE
AUTHORS:
Mukesh Kumar Sharma, Chief Manager, RHQ(T), Indian Oil Corporation Limited, New Delhi
S.G.Venkatesh, Deputy General Manager, RHQ(T), Indian Oil Corporation Limited, New Delhi
1. INTRODUCTION
Grass-root Refinery of 15.0 MMTPA capacity of Indian Oil Corporation Limited
was under construction by IOCL at Paradip. This Project initially envisaged
both Refinery & Petrochemical units. The Petrochemicals envisaged in original
complex consisted of Aromatic Complex producing Para Xylene, Ethyl Benzene,
Styrene Monomer and Polypropylene production from Polypropylene Unit (PP).
Feed for PP unit was envisaged as 93.0 Wt% pure propylene from Propylene
Recovery Unit (PRU) which was part of Indmax project. 93.0 Wt% pure
Propylene is termed as Chemical Grade Propylene. However, during later stage
of the Paradip Refinery project (PDRP), Petrochemical unit’s implementation got
deferred. With deferring of PP unit, disposal problem for Chemical Grade
propylene was envisaged due to lower demand.
Strategy for propylene reduction / disposal of chemical grade propylene
included option like (1) reduction of Propylene Yield from upstream INDMAX
Unit, (2) selling Propylene as LPG that result in Low Value addition, (3) selling
Propylene in the Market as Polymer Grade Propylene (PGP) which is a High
Value Product (99.6 wt % pure Propylene is Polymer Grade Propylene (PGP)).
PDEC (Process Design Engineering Cell), a design arm for Refineries Division of
IOCL, Providing Innovative and Cost Effective Design solutions was referred by
PDRP project group to look into the possibility of converting the Chemical
Grade PRU to produce Polymer grade propylene with minimum modification/
investment. The objective was to produce Polymer Grade Propylene @ 25 T/Hr
(as per demand) as against envisaged production of 150 T/Hr Chemical Grade
Propylene with minimum modification.
This Paper describes how basic fractionation concept with the Combination of
graphical techniques & computer simulation helps in better understanding of
the process and results in optimum design to meet the targeted objective. The
unit has been recently commissioned meeting the desired objective.
1
Category:
Refinery Optimization and Margin Improvement
Paper for RTM 20–22 April 2017, Visakhapatnam, India
2. PROCESS DECRIPTION
Feed to PRU is LPG produced from FCC and Delayed Coker units. The feed
from these units enters into a De-Propaniser column (C-1001) wherein majority
of the C3- are taken out from the top and C4+ material are taken out from the
bottom. C4+ material goes for further processing in downstream alkylation
unit. C3- material goes to Carbonyl Sulfide (COS) Hydrolysis reactor (C-1002)
in which COS is removed. The material is further treated with caustic to
remove H2S and mercaptans. It is then mixed with purge stream coming from
Poly Propylene unit and this mixture is then fed to light end columns (C-1005
& C-1006) in which C2- material is removed from the top and sent to fuel gas.
The bottom material coming from C-1005 containing majority of Propane and
Propylene is fed to Propane/Propylene splitter column (C-1007). Chemical
grade Propylene (93.0 wt% purity) is produced from the top of the column and
sent to Poly Propylene unit. The heavier material containing high amounts of
Propane goes out from the bottom of the splitter column to Refinery LPG pool.
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3. Revised Design Basis
The new design basis for PRU feed and product is as under:
(a) Feed:
Coker LPG
FCC LPG
Total
TMTPA
TMTPA
TMTPA
168.6
954.5
1123.1
Operating Hours Considered 8320
Revised feed composition
(b) Polymer grade Propylene Requirement: 25 MTPH
4. Specification for Polymer Grade Propylene:
It is observed that there is no standard purity specification for Polymer grade
Propylene. In order to fix the design basis for Polymer grade propylene, various
international websites such as Platts were referred.
As per Platts the quotes of polymer grade polypropylene in various regions
are as under:
a. U.S. - Propylene with a minimum purity of 99.0 %.
b. Europe - Propylene with a minimum purity of 99.5 %.
c. Asia - Propylene with a minimum purity of 99.6 %.
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Referring to the above data, the design specification for Polymer grade
Propylene produced from Propylene/Propane splitter column of PDRP was fixed
at 99.6-mole% purity Propylene.
5. Process Modeling of Super-fractionators
Rating of PRU unit with original design feed was performed. De-Propaniser and
Light Ends Columns are normal distillation columns. Any good commercial
simulator can model them with reasonable accuracy.
However C3 Splitter is superfractionator which is characterized by:



Large Number of trays with high reflux ratio and high condenser/reboiler duties
Very High purity of Propylene desired
Components having low relative volatility
Accuracy of Equilibrium constant (K) value is important in superfractionator
systems. For Low Volatility separations, the followings steps are to be carried
out:
 Experimental VLE data / Accurate plant data is required
 Data fit is to be done for VLE using cubic equations of state
 Correct Binary interaction parameters are to be obtained (Interaction
parameter data are normally proprietary)
 These parameters needs to be entered in Process simulator
a. Initial Step - Validation of Existing Design
Interaction parameters were obtained after extensive literature study
and the efficacy of separation was also validated. Reboilers /
condensers design duty as well as other exchanger duty were
validated with existing design. Satisfactory tuning with base data
achieved.
b. Preparation of C3 Splitter Feed
Already validated base model of PRU was used for revised feed. The
comparison of both the feeds is tabulated below:
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Paper for RTM 20–22 April 2017, Visakhapatnam, India
Parameter
Original Design Basis
Revised Design Basis
Feed Rate (T/H)
186
59
Propylene in Feed (% Wt)
79.4
70.8
Propane in Feed (% Wt)
20.4
28.9
Product Propylene (T/h)
147
25
Propylene Purity (% Wt)
93.0
99.7
De-Propaniser and Light End Columns were modeled with revised
feed. These columns were found adequate for revised load.
C3 splitter column rating was started. The Splitter Column has total
of 130 Trays in single column with the feed to 53rd tray from top. The
column is equipped with Vapour Compressor (~13 MW) and design
Reboil duty is 98.0 MMKCal/h. Re-boiler has high flux tubes
(nucleate enhancers). The Splitter column Diagram is as below:
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c. Simulation Observations & Analysis
Simulation of C3 splitter indicated very low quantity of PGP was
possible without any modification in column. Maximum qty of PGP
production possible was10.4 T/hr against requirement is 25 T/hr.
This is mainly due to limitation in reboiler duty of the column.
To produce PGP of 25 T/hr with the existing configuration either
the Column trays (stages) needs to be increased and / Or Re-boiler
Duty needs to be increased. However, this would also call for
augmentation of vapor compressor and re-boiler. Typical C3 splitter
columns producing PGP has around 180-230 trays. But the subject
splitter contains only 130 trays as it was designed for Chemical grade
Propylene.
To produce 25 TPH PGP an option was also there to install an
additional C3 splitter column with around 70-80 trays in series with
the existing column. But this would require additional hardware such
as column, pumps, lines etc & will increase the CAPEX. Additionally
this infrastructure may become infructuous after commissioning of PP
plant.
d. First Principle Approach to debottlenecking PRU to produce
polymer grade propylene
Extensive study done to produce PGP in same column with objective
of minimum hardware modification. Pinch was suspected to be the
bottleneck in design configuration. Tools normally used for detecting /
analyzing column pinch are:
 X-Y Diagrams (Mccabe-Thiele)
 Hengstebeck diagrams
 Key Ratio Plots
i. Key Ratio Plots
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Paper for RTM 20–22 April 2017, Visakhapatnam, India
To identify bottleneck in the column one useful tool is Key ratio
plot. “Key ratio plots” plot key ratio in liquid (XLK / XHK) against
stage number on semi log scale. Slope of the curve provides
measure of relative fractionation accomplished per stage in
various locations along the Fractionator.
Fenske Equation for Min No of stages indicates that
Nmin α ln S
Where,
S = (XLK / XHK)D * (XHK / XLK)B
At total reflux & constant relative volatility log (key ratio) is
linear function of stage number .The curve flattens as the pinch
zone near feed is approached.
ii. Key Ratio Plot for Base Case
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Paper for RTM 20–22 April 2017, Visakhapatnam, India
Analysis of Key Ratio Plot for Base Case
Analysis of Key Ratio plot indicates the following:
 Curve is flat between stages 40 to 70 which indicate Pinch.
 Pinching occurs when operating line approaches equilibrium
curve.
 Pinching wastes stages and sometimes will not accomplish
desired separation.
 Original feed location designed for Chemical grade Propylene
was found pinching for PGP
Column Pinch can be remedied by

Increasing Re-boil and Reflux
o Not Possible in this case due to limitation in Re-boiler
/ Compressor Duty
o Column Diameter also found limiting at required reflux
ratio
o Calls for Major modification (e.g. additional column)

Changing the feed tray location
o In this case alternate feed tray locations are possible
as the column was not yet manufactured.
iii. Study with Alternate Feed Locations
Literature study indicates the following

Key ratio in feed stage liquid should be as close as
possible to key ratio in liquid portion of feed (flashed to
tower pressure)

Feed point should give almost equal slopes on both sides
of feed stage in key ratio plot
Key Ratio in feed liquid for the revised feed is around 2.5. This
key ratio matches with the composition at the lower section of
the column indicating feed tray to be located below the licensor
specified location.
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Key ratio of 2.5 occurs at around stage no. 85. Thus Column
was re-simulated with revised feed tray location. It was
observed that 25 TPH PGP can be produced by relocating feed
tray Column gets de-bottlenecked
Re-boiler duty got
drastically reduced to Only 50% of design duty required to
produce 25 TPH PGP.
The key ratio plots for different feed stage location are shown
below:
Key Ratio plot with Feed Stage 75 (# Tray no 80)
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Key Ratio Plot with Feed Stage 80 (#Tray no 90)
Key Ratio Plot with Feed Stage 85 (#Tray no 100)
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6. Modifications Proposed by IOCL PDEC
The existing Propylene/Splitter column (C-1007) is unable to produce 25 MTPH
polymer grade propylene from revised feed due to limitation in reboiler heat
duty.
To overcome the re-boiler heat duty limitation it is proposed to change the feed
tray location from existing tray no: 53 to tray no: 100 (Tray numbering from top
to bottom) as per the Key ratio plots. This results in reduction of reboiler duty
from around 100 Gcal/h to around 46 Gcal/h and also results in production of
25 MTPH polymer grade propylene or marginally higher from the existing
facility.
It is also suggested that 2 additional feed tray locations at (Tray No: 80 & Tray
No: 90 as numbered from top) may also be put in the column to account for
varying feed quality. This will help in optimizing the production of propylene
Vis-à-vis re-boiler duty during actual operation.
However the existing tray location at 53rd tray was also retained (for future
purpose).
7. Conclusion
Three new alternate feed locations recommended taking care of feed
composition variations. Desired quantity of PGP can be produced with
practically no investment. Additionally no infructuous expenditure shall be
incurred in case of switching to Chemical Grade Propylene in Future. These
recommendations resulted in savings in Capex, Opex and Project
Implementation time.
Thus, the Basic Fractionation concepts helped in de-bottlenecking the system.
Combination of graphical techniques with computer simulation helps in better
understanding of the process and results in optimum design.
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DELTAV EMBEDDED ADVANCED CONTROL
DELTAV EMBEDDED ADVANCED CONTROL
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DELTAV EMBEDDED ADVANCED CONTROL
Table of Contents
1
KEYWORDS _______________________________________________________ 4
2
ABSTRACT ________________________________________________________ 4
3
INTRODUCTION ___________________________________________________ 4
4
MPC CHALLENGES, REQUIREMENTS AND SOLUTIONS __________________ 6
4.1
4.2
5
MPC application tools ................................................................................................ 6
Project Execution ....................................................................................................... 6
EMERSON EMBEDDED APC- EASE OF DESIGN, CONFIGURATION AND
IMPLEMENTATION _________________________________________________ 8
5.1
Plant Operations and MPC sustenance ................................................................... 10
6
EMERSON EMBEDDED APC - HARDWARE, PROJECT SCHEDULE AND
OVERALL COST __________________________________________________ 11
7
EMERSON EMBEDDED APC - IMPROVEMENTS ________________________ 12
7.1
7.2
7.3
7.4
7.5
DeltaV Insight .......................................................................................................... 12
Entech Analyse ........................................................................................................ 14
Entech Tuner ........................................................................................................... 16
DeltaV PredictPro .................................................................................................... 16
Smartprocess® Solutions ......................................................................................... 18
8
CONCLUSION ____________________________________________________ 20
9
REFERENCES ____________________________________________________ 21
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DELTAV EMBEDDED ADVANCED CONTROL
List of Figures
Figure 1
Emerson DeltaV Advanced Control Suite ................................................................. 7
Figure 2
DeltaV InSight: Graphical Interface ........................................................................... 8
Figure 3
DeltaV InSight: Easy launch Options ........................................................................ 9
Figure 4
DeltaV PredictPro: Built-in function block ................................................................... 9
Figure 5
DeltaV PredictPro: Simplified configuration ............................................................. 10
Figure 6
Typical Emerson Advanced Control Architecture .................................................... 11
Figure 7
DeltaV InSight - Enable Learning ............................................................................. 12
Figure 8
DeltaV InSight – Model Analysis .............................................................................. 13
Figure 9
DeltaV InSight – Model Simulation and Robustness Plot ......................................... 13
Figure 10
DeltaV InSight – Adpative Control............................................................................ 14
Figure 11
Entech Toolkit – Power spectrum, Cumulative spectrum, Histogram & Power
spectrum peak plots ........................................................................................................................ 15
Figure 12
DeltaV PredictPro – FIR and ARX models with confidence ...................................... 16
Figure 13
DeltaV PredictPro – Dynamic control objectives for robustness .............................. 17
Figure 14
DeltaV PredictPro – Reference trajectory with funnel and range control .................. 18
Figure 15
DeltaV PredictPro – Online tuning adjustments - Non-linear CV weighting .............. 18
Figure 16
Smartprocess® Distillation – Components ................................................................ 19
List of Tables
Table 3-1: Emerson DeltaV Advanced Control Suite ......................................................................... 4
Table 7-1 Entech Toolkit Analyse - Major Functions ...................................................................... 14
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DELTAV EMBEDDED ADVANCED CONTROL
KEYWORDS
Process Control, PID, Embedded Control, Model Predictive Control (MPC),
Advanced Process Control (APC), Performance monitoring, Emerson, DeltaV
2
ABSTRACT
Numerous instances of using Advanced process control for optimizing plant
throughput, minimizing energy, predicting quality and improving reliability are
available. Advanced process control has also effectively driven process plant to a
new optimum while honouring multiple plant constraints. Justifying and sustaining
Advanced control applications however has numerous challenges some of them
being regulatory control problems, process changes.
DeltaV Advanced Control suite and SmartProcess® applications in addition to the
tools required for model predictive control, quality prediction and constrained
optimization include tools for loop monitoring and adaptive tuning thereby enabling
process control engineers in designing, implementing and sustaining Advanced
control applications including the necessary base regulatory control. Unlike other
control systems with layered advanced applications, DeltaV Advanced Control is
embedded in the system, using the same engineering environment, configuration
database, and controller platform ensuring unprecedented availability and ease of
use.
3
INTRODUCTION
Model Predictive Control (MPC) technology has demonstrated its effectiveness in
dealing with the multivariable constrained problems of the process industry. Over the
years, MPC has developed and evolved with process industry. Significant
advancements have also been made in the tools used for design of MPC as well as
the MPC execution methodology. However, process plants are multivariable. They
have multiple objectives and constraints. The process plants, their variables,
objectives and constraints are all dynamic. Hence, the present challenges and
opportunities involve the ability to design, configure, operate, maintain and sustain
MPC applications with minimum resources, skill and expertise. With the
advancement made in processors and memory it is now possible to achieve this goal
by embedding the MPC in the DCS. This paper provides an overview of products
included with Emerson’s Advanced Control suite which are embedded in Emerson’s
proprietary distributed control system (DCS), DeltaV. Emerson’s product suite
includes tools for application of MPC technology as well as tools for monitoring and
tuning of base regulatory control loops, predicting quality and offline simulation of
process plants. The distinctive advantages of Emerson’s embedded Advanced
solutions are summarized in Table 3-1:
Table 3-1: Emerson DeltaV Advanced Control Suite
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Category
Total
Integration,
Highest
Reliability and
Availability
DELTAV EMBEDDED ADVANCED CONTROL
Solution Benefit
APC Technology resides
in DeltaV Controller
APC Algorithms are
executed in DeltaV
Controller
Quality
and
Productivity
Improved
Security
Sustained
Performance
Detailed Benefits
Controller Redundancy - Unmatched uptime for APC
Data mapping is not required thereby reducing the
points of failure
Common database
External APC server, OPC connection, proprietary
hardware and software are not required
Custom DCS and interface programming, Fail Shed
logic, Watch Dog logic, Data Synchronization is
eliminated
Ease of Implementation, Reduction in Time and Effort
Special training is not
Configuration through Control Studio
required – Can be
Automated Step testing
implemented by Process
Automated Model ID and Validation
Control Engineers
Offline Testing, Simulation and Training
Automated Controller Generation
Automated and Integrated User Interface (Operator
Graphics)
Special software for APC is not required
Security is equivalent to Additional Firewalls, Routers, Switches are not
required thereby reducing hardware and complexity
DeltaV security
External connectivity for APC functions is not required
thereby eliminating security concerns
Emerson is the single focal Point for customer’s
Maintenance
and maintenance needs
Upgrades
APC merges with DeltaV system lifecycle plan
APC upgrades go hand in hand with DeltaV upgrades
and there are minimum compatibility issues during
upgrades
Reduction in Overall Maintenance Cost
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DELTAV EMBEDDED ADVANCED CONTROL
MPC CHALLENGES, REQUIREMENTS AND
SOLUTIONS
The present challenges and requirements from the tools used for design (such as
those for step testing, model identification, controller generation, controller
simulation, optimizer and performance monitoring), project execution, operation,
maintenance and sustenance of MPC applications are well documented.
Some of these are summarized below for reference:
4.1
MPC application tools
 Infinite prediction horizon formulation with reference trajectory, zone and funnel





4.2
control
Nonlinear modeling
Adaptive MPC
Robust stability
Robust identification
Optimizer feasibility
Project Execution
 Identification of regulatory control problems and rectification
 MPC Design - size of MPC
 Project Schedule, cost and budget
 Process changes
 New constraints / Limits
 Change of objectives / Additional objectives
 Model mismatch
 Operator acceptance (Operator reluctance to go out of comfort zone /





Automating operator tasks)
Robustness & performance
Training
Maintenance not budgeted
Online tuning
Benefit estimation
The process plants being inherently dynamic the challenges and requirements which
are addressed during project execution recur during controller operations
necessitating either maintenance of the Advanced control schemes and / or a redesign of the controller which would typically be another project.
The remaining sections of the paper describe the features which enable Emerson’s
embedded DeltaV Advanced control suite to effectively address the above concerns.
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DELTAV EMBEDDED ADVANCED CONTROL
The products which are part of Emerson DeltaV Advanced Control suite are shown in
Figure 1.
Figure 1 Emerson DeltaV Advanced Control Suite
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DELTAV EMBEDDED ADVANCED CONTROL
EMERSON EMBEDDED APC- EASE OF DESIGN,
CONFIGURATION AND IMPLEMENTATION
Significant benefits from APC implementation can be only be achieved by driving the
controlled variable closer to the constraint without violating it. However, to enable this
requires a reduction in variability in the controlled variable which in turn requires a
very stable control at the DCS. Hence, for dynamic control to be effective it needs to
be embedded in the hierarchy of plant control functions.
Emerson’s suite of Advanced control products enable this hierarchy for dynamic
control including tools for monitoring, advanced analysis and stability of the base
regulatory layer. Emerson is the only vendor to provide this functionality within the
control system (DCS). All other vendors provide this functionality outside the DCS in
systems which are less secure, less reliable and non-redundant.
Further, all products are easy to design, implement and use. All Emerson Advanced
control products are graphically configured with pre-built function blocks and built-in
tag browsers. The blocks are developed to encourage process engineers to develop
and deploy APC applications.
The graphical interface and capability to easily launch and use 2 products DeltaV
InSight and DeltaV Predicpro is shown in Figure 2, Figure 3, Figure 4 and Figure 5
below.
Figure 2 DeltaV InSight: Graphical Interface
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DELTAV EMBEDDED ADVANCED CONTROL
Figure 3 DeltaV InSight: Easy launch Options
Figure 4 DeltaV PredictPro: Built-in function block
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DELTAV EMBEDDED ADVANCED CONTROL
Figure 5 DeltaV PredictPro: Simplified configuration
5.1
Plant Operations and MPC sustenance
Typically plant operations are partially involved in the APC project and trained to
operate and maintain APC. However, dedicated involvement of plant operations
during the entire project not only ensures success of the APC but also its sustenance
thereafter. Emerson’s APC products which are easy to configure, implement and reengineer enable fulfilling this need.
The configured APC often requires re-engineering due to inherent dynamics of
process plants such as change in process and APC objectives; increase in model
mismatch, new constraints being identified, to name a few. Emerson’s embedded
tools enable re-engineering with minimum cost and effort ensuring higher uptime of
APC applications.
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DELTAV EMBEDDED ADVANCED CONTROL
EMERSON EMBEDDED APC - HARDWARE,
PROJECT SCHEDULE AND OVERALL COST
Historically, APC projects have required very specialized skill-sets and experienced
resources to implement and maintain – limiting use of technology to very large
refineries or petrochemical plants which could justify this expense. However, unlike
traditional APC systems which are hosted on supervisory computers, Emerson’s
embedded Advanced control can be implemented on the process automation system
(PAS) control network as shown in Figure 6. As part of this implementation APC
functions, can be distributed and executed on existing controllers or Application
Stations. The embedded APC technology significantly reduces the cost of a APC
project and enables justifying APC for all applications which are interactive,
multivariable and can benefit with the MPC technology.
The reduction in overall cost is achieved by reduction in hardware and reduction in
the project schedule because of embedded APC implementation:
 Embedded APC provides a unified interface and communication between the
DCS and APC controller so development of interface for communication is not
required.
 Customized DCS logic is not required for starting / stopping the APC controller
 Embedded APC includes standard graphics for APC operation. So, separate
graphics for APC operation are not required.
 Customized logic is also not required for actions to be performed in the event of
APC controller failure
 Additional databases are not required to be maintained on the supervisory
computer
 The database and controller is the same and so there is no requirement for
time synchronization
 Watchdog timers are not required to confirm whether APC is running or not
Figure 6
Typical Emerson Advanced Control Architecture
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DELTAV EMBEDDED ADVANCED CONTROL
EMERSON EMBEDDED APC - IMPROVEMENTS
Emerson’s embedded Advanced control suite is designed to support implementation,
maintenance and sustenance of Advanced control schemes. A few key features of
the products which enable sustaining plant operations at the optimum are discussed
below:
7.1
DeltaV Insight
7.1.1
Embedded Process Learning, Adaptive tuning and simulation
Embedded process learning when enabled (as shown in Figure 7) calculates
process models and diagnostics for all control loops in the system. The process
dynamics are re-calculated whenever there is a change in set point (or output when
the controller is in manual) and the models updated and stored (as shown in Figure
8). The models are validated using quality parameters for model identification and
model variability. This enables evaluation of process performance over time to
account for process changes and non-linearity.
DeltaV Insight’s Adaptive tuning capability proactively identifies loops that need to be
re-tuned based on the validated models and makes optimal tuning recommendation
using Lamda or Internal Model Control (IMC) tuning rules.
Loop simulation and model analysis (as shown in Figure 9) are included to enable
simulating the loop response and comparing the performance between the existing
and recommended tuning parameters. A robustness plot also enables assessing loop
stability for different tuning parameters.
Figure 7
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DELTAV EMBEDDED ADVANCED CONTROL
Figure 8
DeltaV InSight – Model Analysis
Figure 9 DeltaV InSight – Model Simulation and Robustness Plot
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7.1.2
DELTAV EMBEDDED ADVANCED CONTROL
DeltaV Insight Adaptive Control (Adapt)
The process models and adaptive tuning calculations are used to enable closed loop
control (as shown in Figure 10) to account for changing process dynamics and
process non-linearity. Adapt not only provides optimal tuning parameters but also
remembers the best tuning for recommendation when similar operating conditions
prevail.
Figure 10 DeltaV InSight – Adpative Control
7.2
Entech Analyse
The Entech toolkit enables 2 major functions:
 Entech Analyse can be used to identify the source, frequency and quantum of
variability which can then be eliminated with the Entech Tuner
 Entech Analyse can also be used to identify potential benefits for justifying an
APC project and eventually proving it
The major functions which may be performed by Analyse are compiled in Table 7-1
while few major plots are shown in Figure 11.
Table 7-1 Entech Toolkit Analyse - Major Functions
S.No
Function
Description
Utility
1
Auto
Correlation
The auto correlation is a
measure of the lack of
randomness in the data
It may be used to indicate the absence
of correlation in a feed forward
controlled variable during a load
disturbance
2
Cross
The cross correlation is used to
It enables discovering the coupling of
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S.No
3
4
DELTAV EMBEDDED ADVANCED CONTROL
Function
Description
Utility
Correlation
assess how similar the variability
of one variable is compared to
another.
previously unknown relationships
between variables.
Cumulative
Spectrum
The Cumulative Spectrum is
useful to indicate where there is
significant cycling in a signal.

Determining the amount of
variability caused by a particular
cycle (source).

Deciding which variability sources
should be identified (i.e. payback)
Power
Spectrum
The purpose of the Power
Spectrum plot is to show cycles
in the data.
Identifying common cycles in time series
data.
Identifying sources of variability.
5
Power
Spectrum
Peaks
This is a text plot listing the first
five peaks found in the power
spectrum of the specified
variable
Easily determine period of dominant
cycles.
6
Cross
Correlation
Peaks
The Cross Correlation Peaks
Information window is a text plot
listing the lag and the correlation
coefficient for three peaks
(maximum and minimum) of the
cross correlation plot of the
selected variable against variable
number 1
Easily determine the lag of dominant
correlations.
Figure 11
Entech Toolkit – Power spectrum, Cumulative spectrum,
Histogram & Power spectrum peak plots
The other functions performed by Entech Analyse are time series, statistics, mean,
sigma, 2-sigma, variance, cross-correlation, skewness and kurtosis.
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7.3
DELTAV EMBEDDED ADVANCED CONTROL
Entech Tuner
The Entech Tuner is designed to optimize the tuning of individual loops and the
process using the lambda tuning method which is a method related to Internal Model
control (IMC) and Model Predictive control
The task of customizing the performance of each loop is achieved by specifying how
fast each loop should be by setting the closed loop time constant (called Lambda).
Tuning parameters are determined using the average identified process dynamics.
The tuning calculations are based upon Lambda Tuning methodology and require a
user specified Lambda value along with the specific tuning rule selection.
7.4
DeltaV PredictPro
7.4.1
Model Identification
DeltaV PredictPro algorithm is based on Model Based Predictive Control (MBPC)
and includes a 2-step procedure for model identification to avoid overfitting and
reduce model uncertainty.
In the first step the dead time is calculated over a short horizon using the Finite
impulse response (FIR) followed by complete model identification using the Auto
regressive with external inputs (ARX). This technique ensures identifying an optimal
process model as shown in Figure 12.
Figure 12 DeltaV PredictPro – FIR and ARX models with confidence
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7.4.2
DELTAV EMBEDDED ADVANCED CONTROL
Controller Robustness
DeltaV PredictPro’s robustness is significantly improved by minimizing squared error
of controlled variable over the prediction horizon (prediction horizon dependant
penalty on error) and squared error of controller output over the control horizon
(control horizon dependant penalty on move) as shown in Figure 13.
Figure 13
robustness
DeltaV PredictPro – Dynamic control objectives for
7.4.3
Managing Performance Online
7.4.3.1
Reference Trajectory (Set point filter)
In order to enable managing the controller robustness online, the reference trajectory
(set point filter) with funnel and range control is used as shown in Figure 14. Unlike
typical MPC formulations, instead of penalizing any departure from the trajectory,
only those deviations that are below trajectory or above set point value are penalized
(Area A and Area C if control range = 0). In addition, the control error is considered
zero if the control variable is within range (Area C with range > 0). These features
constitute funnel control, which can be shaped online by changing the setpoint filter
time constant and range, which further enhance controller robustness and flexibility.
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Figure 14
control
7.4.3.2
DELTAV EMBEDDED ADVANCED CONTROL
DeltaV PredictPro – Reference trajectory with funnel and range
Online tuning and Non-linear CV weighting (in MPCPlus)
Additional features (as shown in Figure 15) which enable managing controller
performance online are listed below:
1. Changing the Penalty on error (POE) and Penalty on move (POM) online
2. Non-linear CV weighting – different weights for lower, middle and higher
regions
Figure 15
weighting
7.5
DeltaV PredictPro – Online tuning adjustments - Non-linear CV
Smartprocess® Solutions
Combining domain expertise, consulting services, and Advanced Process Control
(APC) technologies, Emerson industry experts have created a collection of preengineered solutions, collectively termed as SmartProcess®, for optimizing common
process units in DeltaV environment. These solutions serve as a starting template for
optimization projects and are pre-engineered, re-usable, built-for-purpose
configuration templates and control strategies that allow for an accelerated project
schedule thereby reducing time, effort and cost of implementation.
These solutions are presently available for Distillation, Heaters, Fractionators,
Ethylene furnaces. Emerson is in the process of developing many more such
solutions.
Emerson Automation Solutions 2017 – Confidential and Proprietary
05-Apr-2017 – Page 19 of 21
DELTAV EMBEDDED ADVANCED CONTROL
These solutions include (but are not restricted) to the following components as shown
in Figure 16:
 DeltaV PredictPro
 DeltaV Neural
 Standard Calculations & Constraints
 Standard KPI (Key Performance Indices)
 Graphics
 Example configuration
 Optional services for implementation
Figure 16
Emerson Automation Solutions 2017 – Confidential and Proprietary
Smartprocess® Distillation – Components
05-Apr-2017 – Page 20 of 21
8
DELTAV EMBEDDED ADVANCED CONTROL
CONCLUSION
Emerson’s unique embedded Advanced control suite provides the capability within
the control system (DCS) to configure, implement, maintain and sustain optimal plant
operating conditions with minimum resources, time and effort leading to significant
operational and financial benefits. These tools enable the process engineer to quickly
deploy state-of-the-art advanced control to achieve and sustain operational
excellence matches existing industry standards. Emerson provides a user-friendly
interface which is unified and can be totally integrated with the existing control
system ensuring unmatched reliability, availability, security, quality and productivity.
Emerson Automation Solutions 2017 – Confidential and Proprietary
05-Apr-2017 – Page 21 of 21
9
DELTAV EMBEDDED ADVANCED CONTROL
REFERENCES
1.
“A survey of Industrial model predictive control technology” – S.Joe.Qin and
Thomas A Badgwell
2.
“Model Predictive Control: Past, Present and Future” – Manfred Morari and Jay
H Lee
3.
“Easy Robust Optimal Predictive Controller” - Willy K. Wojsznis, Terrence L.
Blevins and Mark Nixon
4. “Embedded APC Tools reduce costs of the technology” – Pete Sharpe and John
Rezabek
5. “Model Predictive Control in Industry: Challenges and Opportunities” – Michael G
Forbes, Rohit S Patwardhan, Hamza Hamadah, R Bhushan Gopaluni
6. “Embedded APC tools dramatically lower implementation costs: A Refinery Case
Study” – Pete Sharpe and Chris Kominar
7. “Advances in Model Predictive Control” – Terry Blevins and Willy Wojsznis
8. “Integration of Real Time Process Optimizer with a Model Predictive function block
“ - Willy Wojsznis , Dirk Thiele, Peter Wojsznis and Ashish Mehta
9. “A Simplified and integrated approach
Implementation”- Vasiliki Tzovla, Ashish Mehta
10. DeltaV BOL (Books-online) V13.3
11. Entech Help Files
Emerson Automation Solutions 2017 – Confidential and Proprietary
to
Model
Predictive
Control
On Site Technical Performance Analysis
of a MicroTurbine
Parivesh Chugh, DGM (R&D)
T.Nandakumar, CM(R&D)
GAIL India Limited
Jubilee Tower, B-35, Sector-1, Noida 201301
ABSTRACT :
Micro-Turbines are energy generators working under the principle of Brayton thermodynamic
cycle with capacity range of 15 to 300 KW. Micro-Turbines have high fuel fungibility and can
run on various fuels like bio-gas, natural gas, propane, diesel, kerosene, methane, and other fuel
sources, thus making them one of versatile equipment for backup power in a variety of
applications. Micro-Turbines use a lean premix combustion system to achieve low emissions
levels at a full operating range. Micro-Turbines have many advantages, including high power
density, light weight, lower emissions, fuel flexibility, low vibration, low maintenance, high
reliability, and excellent durability.
GAIL has implemented a Pilot project at Ghazipur landfill site, Delhi wherein low quality
LandFill Gas (LFG) is being utilized to generate Power using Micro-Turbine technology. This is
a first-of-its-kind project in India where Micro-Turbine is being used to generate 30 KW of
power. This Paper details about the onsite performance of the Micro-Turbine at Ghazipur landfill
site on the parameters of its efficiency, emissions, gross power output and operating hours.
Further, the Paper also explores the possibility of utilization of Micro-Turbine in Oil & Gas
facilities.
Keywords: Micro-Turbine, LandFill Gas (LFG), Efficiency, Emissions
1. BACKGROUND
GAIL has implemented a pilot Landfill Gas (LFG) project at Ghazipur landfill site. The
project is first-of-its-kind in India and implemented over a Landfill area of 10 acres,
adjoining an active Landfill site. The Pilot plant comprises of LFG Extraction and Flaring
which involved closure of allocated landfill site, construction of 20 no’s of LFG extraction
wells, setting-up of enclosed flaring system along with development of associated
infrastructure.
The Project was commissioned in May’2013. Currently, 100 m3/Hr of Landfill Gas with
~25% Methane content is being extracted and safely destroyed in an enclosed flare system
for reduction of GHG and pollution.
As the Methane concentration in LFG is very low, the earlier plan of purifying and upgrading
the LFG to the level of Pipeline quality was not found techno-economically feasible.
Therefore alternate options for utilization of LFG were explored. Based upon an extensive
feasibility work, it was decided to set-up a Micro-Turbine to generate Power after partially
upgrading the LFG.
2. MICRO-TURBINE TECHNOLOGY:
Micro-Turbines are energy generators working under the principle of Brayton
thermodynamic cycle with capacity range of 15 to 300 KW. The Thermodynamic cycle is
indicated in Fig.1
Fig. 1 : Brayton Cycle
As shown above, Micro-Turbines essentially have four main integrated components:
Compressor, Combustion chamber, Turbine and Drive shaft. The Compressors take the
surrounding air at one end of the Micro-Turbine and increase the air’s pressure and density
by compressing it. This air is then fed into the Combustion chamber where it is mixed with
fuel, and then burned. This combustion releases large amount of heat energy and highpressure exhaust gases. The exhaust gases are discharged through exhaust vents into a series
of turbine fan blades that are attached to a central shaft. As the gases are expanded, they spin
the turbine fans, which in turn rotate the drive shaft at high speeds (80,000-100,000
revolutions per minute). The rotational energy produced by the shaft spins the copper coils,
which excite the electrons in the wire, producing electricity. The quantity of electricity
depends on how fast the shaft rotates in the magnetic field, the strength of the magnetic field,
and the quantity and arrangement of the copper coils. To produce electricity at a relatively
low cost, the shaft must rotate at high speed.
Micro-Turbines have high fuel fungibility and can run on various fuels like bio-gas, natural
gas, propane, diesel, kerosene, methane, and other fuel sources, thus making them one of
versatile equipment for producing power. Micro-Turbines use a lean premix combustion
system to achieve low emission levels at a full operating range. Higher air-fuel ratio is used
for lean premix operation and requires operating at high Air : Fuel ratio.
3. ADVANTAGES OF MICRO-TURBINE :
Micro-Turbines have many advantages, including high power density, light-weight, clean
emissions, fuel flexibility, low vibration, low maintenance, high reliability, and excellent
durability. Micro-Turbines can thus be used to operate 24 x 7 as a prime power source.
Micro-Turbines do not require continuous monitoring and thus can operated remotely
making them perfect fit for remote & unmanned locations. Micro-Turbines do not have
lubricants/coolants and thereby no issues w.r.t. Storage, leakage, seepage etc. Micro-Turbines
have high electrical efficiency and heat from exhaust can also be utilized to generate hot
water, steam, chilling, hot air needed for process thereby increasing the overall system
efficiency making them more feasible. Ultra-low emission regulations can also be met by
Micro-Turbine thereby meeting the most stringent of environmental standards.
4. GAIL’s PILOT PROJECT :
GAIL has set-up a Micro-Turbine at Ghazipur landfill site along with a LFG purification
system. This is a first-of-its-kind project in India wherein Micro-Turbine has been installed
to generate 30 KW of power. The Project was intended to demonstrate LFG as a renewable
source of fuel for power generation and to evaluate the Micro-Turbine’s ability to operate
on LFG and to characterize its performance. This Project was commissioned in Dec’ 2015.
The Plant consists of a LFG pretreatment and compression skid, and a 30 KW Capstone
Make Micro-Turbine. The LFG pretreatment and compression skid primarily incorporates
high pressure water pump, absorption tower, intermediate tank, water tank, degasser
chamber, Blower Pump. The skid utilizes a portion of the LFG previously directed to the
flare, and supplies upgraded LFG to run the Micro-Turbine after removal of CO2. The
Process Flow Diagram and the Actual site photograph is shown at Fig. 2 & 3 respectively.
Fig.2 : Process Flow Diagram
Fig.3 : Actual Site Photograph
5. EFFICIENCY OF MICRO-TURBINE:
The efficiency of Micro-Turbine was calculated at various partial loads of output. The Fig.4
shows the relationship between the Power output and Efficiency. The Micro-Turbine has a
recuperater arrangement which utilizes the exhaust heat to preheat the inlet air which
improves the Micro-Turbine efficiency. We were able to achieve > 25% efficiency when the
Micro-Turbine was operated at a load of 20-25 KW. At the full load we are able to achieve
nearly 27% efficiency. The basis of calculations are indicated below :
Input Energy = 111 KW (based upon Calorific value of LFG)
Output power at full load = 30 KW
Efficiency = Output Energy / Input Energy
Observed Electrical Efficiency at Full Load = 30 KW / 111 KW = 27%
In the above calculations, only electrical output is considered and not the thermal output.
30
% Efficiency
25
20
15
10
5
0
0
5
10
15
20
25
30
35
Power Output kW
Fig.4 Power output Vs. Efficiency
The exhaust temperature was also observed against various loads and is shown in Fig. 5.
From the graph it can be inferred that the exhaust temperature is fairly stable in the range of
610-620oF at Load varying from 10 to 30 KW.
700
Exhaust Temp. in oF
600
500
400
300
200
100
0
5
10
15
20
Power output kW
Fig.5. Power output Vs. Exhaust temp.
25
30
Currently, the exhaust heat is not being utilized. However, the exhaust heat can be gainfully
utilized thereby improving the overall efficiency of the Micro-Turbine system.
The total heat potential = m Cp Д T
m = Mass Flow Rate of Exhaust Flue Gas, in Kg/Hr
Cp = Specific Heat of Conductivity of Flue Gas
Д T = Temperature difference between Flue Gas and Exit Temp
Total Heat = 1116 Kg/Hr [Exhaust Flow] x 0.26 Kcal/Kg.oC[Specific Heat] x 140 [Exhaust
delta temp (280-140oC ]
Total Heat = 40824 Kcal / Hr
Two options of waste heat recovery are illustrated as follows:
i)
Cooling Option :
Tons of Refrigeration = Total Heat / 3024 = 40824 / 3024 = 13.5 TR
Thus the system can be used to generate 13.5 TR chilled water
ii)
Hot Water Option :
Increase in Temperature of Water by 20 oC
Total Heat = Water Flow rate [Lit / hr] x 1 [sp. Heat of water] x 20 [Д T]
40824 = Water Flow rate x 1 x 20
Water Flow Rate = 40824 / 20 = 2041 Lit / Hr
So Hot Water Generation = 2041 Lit / Hr
Thus the system can be used to generate 2041 Lit/Hr of hot water at delta 20 degree Celsius.
6. AIR EMISSIONS:
NOx and CO emissions were measured under different loads. The 30 KW Micro-Turbine has
demonstrated a decreasing trend of NOx emissions that stabilized below 10 ppm at 25 KW
which is quite low as compared to a LFG fired reciprocating engine (Fig. 6).
NOx formation is minimized at lower combustion temperature, but lower combustion
temperature results in higher emissions of CO and Total Hydrocarbons (THC). To optimize
low NOx emissions simultaneously with low CO and THC emissions, the combustion of the
fuel must occur at the lowest possible temperature whilst the air and fuel mix must remain in
the combustion chamber long enough to combust most of the fuel. This could be achieved in
Micro-Turbine due to lean premix combustion technology.
140
120
100
Nox Emission in ppm
80
60
40
20
0
0
5
10
15
20
25
30
Power Out Put kW
Fig.6. Power output Vs NOx
CO is a criteria pollutant because of its potential detrimental impact on human wellbeing. It
is a poisonous gas formed when carbon based fuel is not fully burned. The CO emissions
were also measured at various loads and CO formation also decreased continuously as the
load was increased. It is less than 2 ppm with load > 15kW (Fig. 7). It shows that LFG has
been completely combusted in the Micro-Turbine.
16
CO emmissions in ppm
14
12
10
8
6
4
2
0
0
5
10
15
20
25
30
Power output KW
Fig.7 Power Output Vs.CO emissions
7. GROSS POWER OUTPUT:
The Micro-Turbine power output is directly related to ambient air temperature. At 200m
elevation and ambient temperatures above 35° C, the predicted maximum power of the
Micro-Turbine in Delhi is less than 30 KW. The predicted maximum power output of the
Micro-Turbine increases as the temperature drops below 35° C.
The Micro-Turbine efficiency with respect to ambient temperature was measured and is
shown in Fig-8. This data is from December to February months which have typically low
temperatures. The maximum power produced by the Micro-Turbine during the 3-month
period was 28 KW at 15°C, i.e. 93 % of nameplate rating. Overall, the power output did
correspond to the predicted maximum power curve.
Power output & Efficiency
35
30
25
20
15
10
5
0
Dec
Jan
Feb
Power Output
Mar
April
Efiiciency
Fig.8 A : Power Output & Efficiency Vs. Trial Month
Ambient Temp & Efficiency
35
30
25
20
15
10
5
0
Dec
Jan
Feb
Ambient Temp
Mar
April
Efiiciency
Fig.8 B : Ambient Temp. & Efficiency Vs. Trial Month
8. MICRO-TURBINE OPERATIONS:
During the 3-month demonstration period, the Micro-Turbine was run on an average
of 8 hrs/day. The Micro-Turbine is connected and synchronized with the Grid Power.
Therefore its performance is dependent on Grid Power quality. The LFG site is plagued with
wide power fluctuations due to power theft in the surrounding slum areas. Although voltage
stabilizer was installed, still the wide fluctuations in voltage led to tripping. Moreover, any
Grid failure also leads to tripping of Micro-Turbine.
The main factors that contributed to frequent tripping of Micro-Turbine were :




Grid power quality issues
Oxygen ingress in LFG
Variations in LFG quality & quality
General Maintenance & Modification works
Any tripping of Micro-Turbine leads to lot of down time in restarting the operations and in
this Case the tripping were frequent. Sometimes the downtime even lasted upto 12 hours due
to multiplicity of factors.
9. OTHER LFG CHARACTERISTICS:
a) Siloxane concentration:
The Siloxane in LFG should be <5 ppb for safe operations of Micro-Turbine as it forms gum
like substance in the combustion nozzles. Siloxane values obtained in the LFG feed stream
for the 3 months are shown in Table No. 1 :
Month
Siloxane
concentration
December
Non-traceable
January
-
do -
February
-
do -
Table -1 : Siloxane concentration in LFG
b) H2S concentration:
The Hydrogen Sulphide (H2S) in LFG should be < 4000 ppm for safe operations of MicroTurbine. The H2S values in the LFG feed stream for the 3 months are shown in Table No. 2 :
Month
Hydrogen Sulfide
December
< 5 ppm
January
< 5 ppm
February
< 5 ppm
Table -2 : H2S concentration in LFG
10. SUMMARY:
The above performance analysis indicates that Micro-Turbine is a versatile system having
high turn-down ratio that can be used to generate power from even low quality fuel like
LFG. The performance w.r.t. efficiency and emission standard is also meeting the
benchmarks with high turn-down ratio.
There are various proven advantages that make Micro-Turbines appealing. From an
economic standpoint, the Micro-Turbines have few moving parts and consequently are easy
to operate and maintain in comparison to conventional gas or diesel or Fuel cell power
plants. Another advantage of Micro-Turbines is their durability and reliability. MicroTurbines also have high Power density as these produce a large amount of energy relative to
their size with low emissions.
Micro-Turbine technology can be used in various facets of Oil & Gas business to improve
efficiency. Few of the possible uses could be to generate power in remote locations and
isolated fields, Flare Gas recovery etc.
BIBILIOGRAPHY:
i.
Pilavachi, P. A. "Mini-and micro-gas turbines for combined heat and power"
Applied Thermal Engineering 22.18 (2002): 2003-2014.
ii.
Basrawi, Firdaus, et al. "Effect of ambient temperature on the performance of micro
gas turbine with cogeneration system in cold region." Applied thermal
engineering 31.6 (2011): 1058-1067.
iii.
Moya, M., et al. "Performance analysis of a trigeneration system based on a micro
gas turbine and an air-cooled, indirect fired, ammonia–water absorption
chiller." Applied Energy 88.12 (2011): 4424-4440.
iv.
Kim, Sunhee, Taehong Sung, and Kyung Chun Kim. "Thermodynamic Performance
Analysis of a Biogas-Fuelled Micro-Gas Turbine with a Bottoming Organic Rankine
Cycle for Sewage Sludge and Food Waste Treatment Plants." Energies 10.3 (2017):
275.
v.
Pantaleo, Antonio M., et al. "Energy performance and thermo-economic assessment
of a Micro-Turbine-based dual-fuel gas-biomass trigeneration system." (2016).
Holistic Refinery well-being through efficient and robust Amine system
Ashok Kumar, Chief. Manager-Technical; Ashok P. Golekar, Sr. Manager-Technical;
Satish D. Khedekar Manager – Technical; Gaurav Vyas Sr. Engineer - Technical
Mumbai Refinery, Hindustan Petroleum Corporation Limited
INTRODUCTION
Amine units are an integral part of any refinery and are an important tool for the removal of
acid gas from petroleum products and intermediates. The operation of amine units is often
overlooked because it is not seen as revenue center. Overlooking the amine treating system can
lead refinery to many shut downs, outages and not meeting product specifications or emission
norms. However In recent years, across all refineries there has been increased emphasis on
improving amine system efficiency due to the increased processing of sour crude and higher
throughput than design, especially in HPCL Mumbai Refinery which is processing close to 8.2
MMTPA against design of 6.5 MMTPA . This triggers the development of several optimization
ideas to match up an effective and robust amine system with the additional refinery processing
capacity.
After the turnaround of May 2015, wherein we found several equipment in the amine system
severely corroded, coupled with past history of amine system problems of high apparent H2S in
lean amine and hydrocarbon carryover to SRU, a need for overhaul of amine system was
necessary. Turnaround of the refinery May 2015 presented a great opportunity to do a fresh
start. After startup of ARU unit, simple plant operational procedures changed and maintenance
steps were taken to improve amine system. After the teething troubles of around 4 months,
these steps not only also helped in effective removal of acid gases but also created a robust and
reliable amine system which aided to holistic refinery well-being and profitability.
Amine system improvement was easier said than done as it is spread across various operating
units and facilities, problem in any of the operating area can trigger problem in all other area
and amine regeneration, so control of entire circuit was not was not so straightforward.
Actions were taken in a systematic and methodical way to gauge the impact of changes made.
AMINE SYSTEM IMPROVEMENT AND REFINERY WEL BEING; CASE STUDY
1.0 PROBLEMS FACED:
Problems encountered before Amine system improvement steps were taken up:
 Excessive foaming in regenerator and absorbers
 Rapid fouling of lean/rich exchanger
 Upsets in FCCU area contactors and carryover of amine to fuel gas
 LPG failing on copper strip corrosion test
 High SOx in furnace flue
 Lean amine Cartridge filter high pressure drop
 High apparent H2S in lean amine
 Hydrocarbon in acid gas and upset in SRU
 High HC flare
2.0 OPERATIONAL STEPS IN CREATING A ROBUST AMINE SYSTEM:
Operation at optimum conditions is the first key to enhance the performance of the amine
system. Optimum performance of the amine sweetening unit depends on proper tuning of the
operating conditions.
a)Circulation rate and Concentration:
Lower amine concentrations are generally considered to be precaution against corrosion.
However, low amine concentrations, in combination with under-circulation of amine, lead to
higher rich acid gas loading, thereby causing severe corrosion in the rich amine line. Also, low
amine concentrations require a higher amine circulation rate and/or a higher reboiler steam
rate to achieve sweet gas specifications in terms of acid gas removal.
Amine system was operating at higher circulation rates. Higher circulation rates needed more
energy in pumping, created more losses, needed more steam in stripper, increased tower liquid
velocities, increased foaming tendency and increased entrainment rates. Thus to reduce amine
circulation rate concentration of MDEA was increased from 20-25% levels to 35% level.
Why circulate more water in your amine system than necessary just to heat and cool it
down? High Concentration provides a significant reduction in solvent circulation rate, steam
consumption rate, pumping duty and unit load while posing less risk of hydrocarbon carryover,
less corrosion due to lesser velocity.
With concentration and circulation changes, rich amine H2S loading was increased from an avg
of around 0.3-0.33 mol/mol to around 0.42-0.45 mol/mol
Parameters
Prior to Improvements
Post Improvements
(Oct 2015)
Amine Concentration
20-25%
35%
Circulation rate, m3/hr
230-240
165-215 m3/hr
Amine H2S loading
mol/mol
0.3-0.33
0.42-0.45
Lean amine supply
temperature (Deg. C)
45-46
48-50
b)Routing of Rich amine flash drum off gases to nearby unit furnace
Rich amine flash drum off gases which are rich in propylene and ethylene were sent to near-by
unit (Prime G) heater furnace. Utilization of rich amine flash drum off gases has led to around
25-30 MT of fuel gas saving every month and this scheme has potential to generate saving of
around 3.3 Crore per annum. By using ARU off gas to Prime G furnace not only off gas as energy
source was utilized but also it helped in holistic refinery objectives of cutting carbon emissions.
Sour flare mitigation – before commissioning of the scheme average sour gas flow in last 120 days of
commissioning was 848 kg/hr which reduced to 249 kg/hr post commissioning (20 Aug to 19 Sept 2016)
period.
Graph Showing Sour flare load before and after 60 days of commissioning.
Sour flare @ 120d0h0m 0s
4000 6/23/2016 14:36:36
10/21/2016 14:36:36
3000
2000
1000
0
** #1 (R) 509fi5202.daca.pv 10/21/2016 14:36:36
-18315.142970562 K (Avg @1 Day)SOUR FLARE GAS FLOW
(X axis- dates 20 June to 19 sept & Y-axis sour flare flow in Kg/hr)
c) Activated carbon & slip stream circulation
Activated carbon was earlier coconut based charcoal which was changed to coal based
activated carbon, same has resulted in effective removal of HC from the amine and also
increased charcoal media life as compared to coconut based filter was found. Charcoal based
filter media lasted 14 months and it was found to be working in good working condition then
also.
d) Slip stream circulation:
Slip stream circulation rate was maintained at 10% of lean amine flow. So as to contain
corrosion particles in the system, which can initiate foaming by providing surface for the
foaming. Due to high apparent H2S in lean amine in past there was possibility of lots of
corrosion particles in the system, for which cartridge filter media was replaced and specification
was given 5 micron. This along with several other changes resulted in foaming free system.
e)∆ T betwee supply lea a i e a d treati g gas
Proper amine cooling enables optimal rich loading as amine absorbers work on temperature
and concentration gradients. Cooler lean amine temperatures maximize the mass transfer of
H2S from process gas to the amine solution. However minimum temperature differential is
required to prevent hydrocarbons carryover into the amine solution.
Lean amine supply temperature was kept 48-50 Deg C so as to maintain 10oC ∆T between lean
amine and treated gas. Maintaining of this temperature differential prevented hydrocarbon
condensation in the lean amine.
f)Foaming control
Foaming is caused by particulate built up (corrosion and other particulates) and hydrocarbon
carryover in rich amine. Foaming was controlled by antifoam agent, HC carryover control, HC
built up prevention through slip stream, and by increasing concentration of amine and reducing
flowrates in absorber and regenerator columns.
g)Change of Antifoam
Antifoam dosing of glycol based antifoaming agent was started to control foaming, dosing rate
was kept 3 to 7 ppm in lean amine.
Control of foaming reduced amine losses in absorbers such as FCCU absorbers where carryover
of amine to fuel gas drum and subsequent amine loss was common. Addressing foaming
problem resolved problem of high apparent H2S in amine, LPG and Fuel gas quality and amine
loss as well.
h)Anticorrosion agent dosing
Not only effective removal of the corrosion particles but prevention of corrosion is also
imperative for a robust and effective amine system. Anti-corrosion agent dosing program was
employed to create a corrosion free system. Iron particles are measured in Reflux drum every
fortnight to check corrosion inhibitor effectiveness.
Corrosion in amine treating units is a well-known phenomenon that can result in down time
and lost production.
i)Heat stable salts
Heat stable salts can tie up an amine molecule and reduce active amine available for acid gas
removal. Heat stable salts also facilitate corrosion reactions and act as chelating agent.
Sample for heat stable salts and possible mitigation was undertaken with various agencies
working in the domain and samples were sent regularly to check heat stable salts, anions and
cations in the lean amine. In our refinery heat stable salts in Lean were found 0.59 wt% which is
quite less than 1.5% max. value, therefore any remedial treatment for heat stable salts removal
was not envisaged. Though agencies for the HSS removal were kept in contact for HSS removal
services offered.
j)Nitrogen supply in Amine storage tank
Amine color was checked regularly, it was found brownish, indicating amine degradation
because of oxygen. Moreover in anions and cations analysis of lean amine – acetate and
formate were found high indicating degradation due to oxygen. Phosphate was not high
therefore corrosion inhibitor rates were not reduced. Nitrogen supply in amine storage tank
was ensured to prevent oxygen contamination at all times. Oxygen contaminated amine further
leads to high corrosion and generation of corrosion particles.
Particulars
Sample October 2015
Sample Feb 2016
HSAS, wt%
0.7964
0.6986
Inorganic HSS
0.5611
0.1206
Total HSS
1.35
0.8192
Formate (ppm)
1082
954
Acetate (ppm)
3789
1389
Thiosulphate
(ppm)
2497
357
#Results given by Amines & Plastizers lab, Turbhe, Navi Mumbai
k)Monitoring of Amine
Lean Amine color was regularly monitored to check contamination, presence of hydrocarbon,
H2S. Shake test shaki g a i e igorously i a ottle as do e e ery eek to he k the
foaming tendency of amine and based on that antifoam dosing rates were adjusted. Even a
good amine can form a foam layer but this should not be persistent (less than 2 seconds). If the
amine color appears pitch black amine (as was case with our plant before improvements),
system improvement can take several weeks. When foam takes a longer time to break then it
indicates amine solution is starting to accumulate contaminants and troubles in amine system
may start. Regular monitoring of amine system is required to keep amine system problem free.
l)Filter media change frequency
6 months replacement was fixed for cartridge filter and 1 year for carbon filter to keep amine
system in a healthy state.
m)Rich amine flash drum Hydrocarbon side skimming frequency
Skimming frequency was fixed to once per day to remove Hydro carbon to rich amine flash
drum Hydrocarbon compartment.
n) Amine Regenerator Reflux drum water draining to SWSU
This water from ARU reflux drum is sent to SWSU once in a day to prevent accumulation of heat
stable salts in lean amine.
o) Stripping Ratio
Only enough steam into the tower to be sent to reach a lean amine loading target. We
observed that 0.065 MT steam/MT of feed was good enough to get H2S stripped after the
changes were made. Design of the unit is 0.1 MT steam MT/MT of feed. By reducing circulation
rate and concentration 35% saving in steam compared to design was observed.
p) Nickel passivator addition in FCCU
This was added in FCC feed to reduce H2 content in off gas generated, it reduced H2 content in
off gas around 15%, which in turn affected volume flow of FCC fuel gas making amine absorber
adequate by reducing vapor liquid load.
q)Emulsion formation control on FCC side
Demisters pads installed in off gas KOD prior to absorber to control heavier hydrocarbon
ingress in absorber. Stable emulsion used to form in the absorber column leading to amine
losses and loss of tower operating control and LPG going off-spec on H2S. Emulsion breaker was
employed in FCC side. This reduced amine losses and carryover of amine.
CONCLUSION AND COST BENEFIT ANALYSIS:
It is easy to identify the cost added in amine system - amine replacement, filters, activated
carbon however larger impact on product going off-spec due to amine system performance is
often discounted. Amine unit do not produce saleable products, but this does not mean that
optimization of this cannot yield large benefits. There are many benefits that can be realized by
an effective and robust amine system such as consistent product quality, less energy cost,
improved operation of sulphur recovery plant, emission under control.
Benefits attributable to Amine System
improvement
Monetized value per annum
Cost saving attributable to consistent onspec product
Rs. 19 Crores ( 2.92 Million USD)
Saving due to reduction in steam
consumption
Rs. 8.2 Crores ( 1.26 Million USD)
Savings by reduction of circulation rate
Rs. 0.55 Crores ( 0.085 Million USD)
Amine loss reduction
Rs 0.24 Crores ( 0.037 USD)
Rich amine off gas to nearby heater
furnace
Rs 3.28 Crores ( 0.5 Million USD)
Sox Emmissions/Fuel gas
Intangible (Environmental)
Sour flare Reduction
Intangible (Environmental)
Total per annum
Rs. 31.3 Crores ( 4.81 Million USD)
In the case study, there were several areas that benefited from amine system improvement
however the one most benefitted in terms of monetary value was LPG amine absorber. One
LPG bullet used to go offspec every month before amine system improvements were taken up.
If this cost is considered this is equal to saving of approx. Rs. 19 Crore/annum
Amine concentration increase and other improvements lead to reduction in steam reduction to
0.065 steam MT/MT feed lead to saving of Rs. around 8.2 crores/annum.
Savings by reduction of circulation rate is around 0.55 Crores and savings in form of reduced
amine loss is 0.24 crores. ARU off gas to Prime G unit has potential of saving Rs 3.28
crore/annum.
Thus it can be concluded that amine system improvement if monetized has aided benefit of
around Rs. 31 crore to the refinery bottom line, moreover it has aided in several intangible
benefits such meeting refinery objectives of sour flare mitigation refinery and energy
consumption reduction.
Effect of ZSM-5 zeolite crystal size in Fluid Catalytic Cracking Reactions
Praveen Chinthala, Gopal Ravichandran and Asit K Das
Refining R&D, Reliance Technology Group
Reliance Industries Limited, Motikhavdi, Jamnagar, Gujarat – 361142
Abstract
Fluid catalytic cracking (FCC) is one of the largest secondary refinery process in which
VGO/ HTVGO is converted to value added products namely propylene, LPG, gasoline and
middle distillates.. FCC catalyst and ZSM-5 additives are being used in FCC process to obtain
the desired product yields; wherein the former increases the conversion by primary cracking and
later maximizes propylene and butylene by secondary cracking. Hydrothermally stable ZSM-5
additive is essential for maximization of lower olefins (=C2 to =C4) and we have been able to
develop the hydrothermally stable additives for these applications. The optimum crystal size of
zeolite provides good activity/selectivity and to have good hydrothermal stability in gasoline
cracking. In the present study, three ZSM-5 zeolites having different crystal size small, medium
and large ranging from 50-1000 nm were formulated in the additive formulations.
The catalytic performance of the additive was then evaluate by cracking HT-VGO feed
with USY based commercial FCC catalyst. The performance results shown that lower crystal
size zeolite (<300 nm) containing formulations are superior in yielding higher propylene by
gasoline cracking in comparison to higher crystal size (1000 nm) zeolite containing formulations.
It is found that our formulation has provided good hydrothermal stability for lower crystal zeolites
and easy accessibility of active sites, thereby to achieve better cracking activity for maximization
of propylene in FCC.
1.0 Introduction
Conventional processes for catalytic cracking of heavy hydrocarbon feedstock to
gasoline and distillate fractions typically use a large pore molecular sieve, such as zeolite Y, as
the primary cracking component. It is also well known to add medium pore size zeolite, such as
ZSM-5, to the cracking catalyst composition to maximize lower olefins and/or increase in the
octane number of the gasoline fraction. Hydrothermal stability of zeolite in the additive catalyst
is one of the major concerns in FCC in which the catalyst and additive are subjected to higher
temperatures in presence of steam in a regenerator. This results in deactivation of
catalyst/additive due to dealumination of zeolites in the catalyst/additives. Thermal deactivation
1
of FCC catalyst is not a significant in comparison to hydrothermal de-activations. Generally FCC
catalysts are hydrothermally deactivated at 750 to 800 ºC in the laboratory/ pilot plant for 5-20
hrs to simulate commercial FCC plant yields. The matrix (silica, alumina, clay) component gets
slow deactivation at high temperature of 700-800 ºC over the time, however, no appreciable
deactivation takes place under hydrothermal conditions. On the other hand, zeolite is
susceptible to hydrothermal treatment due to cleavage of Si-O-Al bonds.
The zeolite
dealumination takes place with or without loss of surface area, which is much depends on the
zeolite type, method of preparation and experimental deactivation conditions. The phosphorous
play a key role in stabilization of ZSM-5 zeolite, thereby minimizing the dealumination in FCC
condition.
It is well known in the prior art [1] that conventional ZSM-5 cracking additives have a
crystal size > 0.2 microns since smaller or nano crystal zeolites have lower hydrothermal
stability and hence rapidly lose activity when exposed to high temperature steam generated
during FCC regeneration. This is a concern in conventional FCC catalyst technologies to
develop hydrothermally stable nano zeolites and incorporate in FCC catalyst/additive
formulations. Hence, preparing hydrothermally stable FCC catalyst additive with nano zeolites is
one of the challenges and the present study addresses the stabilization of nano zeolites for FCC
applications.
Maximizing propylene from FCC is a prime objective of Reliance as it provides max
value to our refinery and hence it has high business relevance in poly propylene. RIL is focused
on R&D activities in the Refining sector with an overall objective to create value through
innovative solutions and cutting edge technology developments. Refining R&D, in a short period
has been able to develop few breakthrough technologies related to propylene maximization in
FCC including catalyst and processes [2-4].
Thus, in the current study, it has now been found that the addition of phosphorous
containing medium pore ZSM-5 zeolite having a crystal size less than 300 nm incorporated in
additive formulation. These additives added to a conventional large pore molecular sieve
cracking catalyst in FCC evaluations and which improved propylene selectivity in the catalytic
cracking of VGO without loss in activity.
2.0 Experimental
2.1 FCC catalyst additive preparation
2
The primary base FCC catalyst is commercial catalyst. The catalyst additive (ZSM-5) is
prepared by spray drying an aqueous slurry containing different crystal sizes of zeolites as
shown in Table 2. The slurry was prepared by homogeneous mixing of individual slurries
contains about 40% ZSM-5 zeolite, clay slurry, binders (silica, alumina) and zeolite stabilization
of phosphorous as per our patented procedure [3]. The spray dried product is subjected to
calcination at 500 ºC for 1h prior to its hydrothermal deactivation.
2.2 Fluid Catalytic Cracking Reaction
Advance Cracking Evaluation (ACE) unit is a fixed fluid bed reactor used for the
evaluation of catalyst additives of the current study. The commercial base catalyst and different
crystallite size ZSM-5 zeolites containing additives were hydrothermally deactivated at 800 ºC
for 20 and 100 hrs respectively prior to the performance evaluation. The catalyst and additive
are used in a mixture of 75% catalyst and 25% additive for the evaluation. The reaction was
carried out at four different catalyst-to-oil ratios (4 to 10) to generate wide range of conversion
data. The reaction temperature was maintained at 545 °C. Product gas and liquid was analyzed
in Agilent 3000A micro GC and Varian 450 GC SIMDIST respectively.
Table 1: ACE operating conditions
Parameter
ACE protocol
Feed injection time
Fixed at 30 seconds
C/O range
4 to 10
Feed Rate, gm/min
2.0
Rx temp. °C
545
Rx. Pressure, kg/cm2(g)
Atmospheric operation
Results and discussion
The physical properties of 3 proprietary zeolites of different crystal size are analyzed by physical
properties are compiled in below Table 2.
Table 2: surface area and pore volume data of different crystal size zeolites
Zeolite physical
properties
Total surface area, m2/g
SiO2/Al2O3
Z1*
Small crystal
380
24
Z2*
Medium crystal
355
25
3
Z3*
Large crystal
376
29
SEM crystal size, nm
60
150
1000
* Proprietary zeolites for FCC applications.
b) ZSM-5 (Z2) – 150 nm
a) ZSM-5 (Z1) – 60 nm
c) ZSM-5 (Z3) – 1000 nm (SAR
Figure 1: SEM micrographs of ZSM-5 zeolites with different crystal size
Figure 1 shows the FE SEM micrographs of the zeolites with different crystal size having SAR in
the range of 24-30. Three zeolites of different crystal size are used in the current study and their
observed crystal sizes were 60 nm, 150 nm, 1000 nm and they have named as Z1, Z2 and Z3
respectively.
The spray dried ZSM-5 additives after calcination are analyzed by various physico-chemical
properties are summarized in below Table 3.
Table 3: Physico-Chemical Properties of prepared ZSM-5 additives containing different crystal
size zeolites
Catalyst additive
Additive-1
4
Additive-2
Additive-3
Physical properties of calcined add samples
TSA (F), m2/g
117
112
123
ZSA (F), m2/g
62
67
87
2
MSA (F), m /g
55
45
36
TPV, cc/g
0.17
0.14
0.09
ABD, g/cc
0.79
0.74
0.77
APS, (µ)
90
85
95
Attrition Index (AI)
3
7
11
%Crystallinity
27
36
33
Surface area of hydrothermally deactivated additive
samples (100 hrs)
TSA(S), m2/g
141
148
154
ZSA(S), m2/g
66
58
79
2
MSA(S), m /g
75
90
75
The ZSM-5 additives 1, 2 and 3 are prepared from clay slurry, binders (silica, alumina), and
phosphorous stabilized ZSM-5 zeolite slurry. All the raw material used in the formulation are
same except the crystal size of ZSM-5 zeolite. 60 nm, 150 nm and 1000 nm crystal size ZSM-5
zeolite containing formulations are named as Additive-1, Additive-2 and Additive-3 respectively.
The matrix surface area and total pore volume are increasing with decrease in crystal size of the
zeolite.
Table 4: FCC product yields at 76% conversion
FCC product
yields wt%
Base
catalyst
+Additive-1
Base
catalyst
+Additive-2
Base
catalyst
+Additive-3
C/O ratio
5.2
5.3
5.6
Coke
4.1
4.3
4.1
Fuel gas
2.1
2.0
2.1
Propylene
13.9
13.5
12.6
Gasoline
35.6
36.5
38.8
LCO
16.0
15.8
16.2
CSO
8.02
8.2
7.8
Total LPG
34.3
33.2
31.8
5
The performance evaluation results found to show that lower crystal zeolites (<200 nm)
containing additives 1 and 2 have shown improved propylene yields. 0.9% higher propylene
yield in Additive-2 and 1.3% higher propylene yield in Additive-1 are observed due to the
presence of nano crystal zeolites. These superior results achieved without reduction in activity
and no increase in coke and fuel gas. It may be noted that the improved performance was after
severe hydrothermal deactivations of additives which is significant improvement and excellent
hydrothermal stability. The nano zeolites (<0.3 μ) facilitates to overcome the diffusion limitations
and thereby enables more number and the accessibility of acid sites for cracking. Thus, the
higher cracking of gasoline range hydrocarbons led to yield higher propylene yield in FCC from
nano zeolites containing additive formulations in comparison to larger crystal zeolites (1000 nm)
containing formulations.
The proprietary nano zeolites and their effective stabilization by phosphates is being
used. The direct method effectively used the phosphorous to stabilize the zeolite by aging and
direct contacting the zeolite with minimum clay-phosphate interactions during preparation.
Further, the synergic effect of silica/alumina binders, clay, and zeolite-phosphate interactions
led to better zeolite stability.
Summary
The additives of the current study with different crystal size of ZSM-5 zeolites from 50 nm to
1000 nm were examined in fluid catalytic cracking along with base catalyst. The lower crystal
zeolite (<0.3µ) containing formulations exhibited superior performance in yielding higher
propylene yields (>1%) in FCC. The nano zeolites containing additive in our proprietary
formulation enable to have stability even under severe hydrothermal deactivation conditions at
800 ºC for 10-100 hrs. The proprietary nano zeolites and their stabilization technology
developed by us are the key factors responsible for its superior performance.
Acknowledgements
The authors acknowledge the RIL management for permitting to publish this work. The authors
are also thankful to Dr. Vijay B for analyzing SEM and FCC Process & Catalyst teams specially
Sudhir, Mehul and Vinodh for experimental support.
References
[1]
[2]
[3]
[4]
US patent no. 4828679
Ravichandran et al. US patent no. 9067196
Chinthala et al. WO 2016/087956 A1
Mandal et al. US9550708B2
6
Biography
Praveen Chinthala is a General Manager in the Refining R&D Division of Reliance Industries
Limited (RIL). He has 17 years of experience in FCC and heterogeneous catalysis with 9
patents and 17 publications to his credit. He holds degrees in Chemistry (MSc and PhD) from
Osmania University and IICT, Hyderabad respectively.
E-mail: [email protected]
Gopal Ravichandran is Assistant Vice President, Lead FCC catalyst group in Refining R&D
Division of RIL. He has 23 years of industrial experience in Refining catalysis with >20 patents
and >10 publications to his credit. He holds degrees in Chemistry (MSc and PhD) from IIT
Mumbai and engineering (MTech) from IIT, Kharagpur respectively.
E-mail: [email protected]
Asit Kumar Das heads the refining R&D division at RIL, Jamnagar. He has 31 years of
experience in Refining Research with >40 patents, >40 publications and several book chapters
to his credit. He holds degrees in chemical engineering (BTech, MTech and PhD from Jadhapur
University, IIT Kanpur, and University of Gent, Belgium, respectively.
E-mail: [email protected]
7
Recovery of hydrogen from the refinery off-gas streams: Application of membrane
technology
Nitin Somkuwar, Renny Andrew, Sonal Maheshwari, Gokak D.T.
Corporate R&D Centre, Bharat Petroleum Corporation Ltd., Greater Noida, U.P.
Email: [email protected]
In Refinery, hydrogen (H2) gas is the most essential and expensive gas, which is mainly used for
hydroprocessing in refineries. The refineries have various streams rich in hydrogen and which are routed
into the fuel gas for burning in furnaces due to unavailability of economic recovery process. The
hydrogen concentration in the streams typically varies from 20-80 vol%.
To date, large scale hydrogen production generally occurs via stream methane reforming (SMR)
followed by Water Gas Shift (WGS) reaction. High purity H2 is generally produced using Pressure Swing
Adsorption (PSA). Generally, PSA works at 80-85% recovery and leftover H2 goes as off gas to fuel gas
header. However recovery of leftover hydrogen from the different streams in refinery would add value
by reducing the hydrogen consumption in refinery.
Hydrogen enrichment can be achieved by various approaches including pressure swing adsorption (PSA),
cryogenic distillation (CD) and membrane separation (MS). Whereas, PSA and CD are highly energy
intensive processes compared to membrane separation. Typically, membrane separation exhibits the
following characteristics; higher energy efficiency, cost effectiveness, simplicity in operation,
compactness, portability, longer life, better separation at ambient conditions and better compatibility
with environment.
The paper highlights mainly the comparison of available membrane technology options, techno
commercial studies and indigenous development of polymeric membranes.
1. Introduction
In Refinery, hydrogen (H2) gas is an important and costlier reactant, used for mainly hydrocracking and
hydrotreating processes. Today, the most widely used method (80 %) for H2 production is the steam
reforming of light hydrocarbons. High-purity H2 is generally produced from many refinery process units
including HGU, HCU, CCR etc. Presently huge amount of hydrogen is left out in to the fuel gas stream
due to non availability of economically viable recovering processes. However recovery of leftover
hydrogen from the different streams in refinery would add value and reduce the load on hydrogen
production units.
2. Hydrogen purification technologies
The purity and pressure of the hydrogen stream available to in refinery units have significant effect on
the design and operating of processing units. The three main hydrogen purification technologies used in
refineries are pressure swing adsorption (PSA), selective permeation using membranes, and cryogenic
separation. Each of these options based on a different separation principle, and consequently, the
characteristics of these processes differ significantly. The appropriate hydrogen purification technology
selection, depends not only on the economics, but also, on flexibility, reliability, and easy of future
process expansion.
2.1. Pressure swing adsorption (PSA)
PSA is a hydrogen purification process in which the impurities consist of CH4, CO2, CO, H2O, etc. in a gas
stream are removed in adsorbent beds. PSA units are based on the ability of adsorbents to adsorb more
impurities at high gas-phase partial pressure than at low partial pressure. The adsorbents, depending on
specific application, are usually made of molecular sieve, activated carbon, activated alumina or silica
gel. In this process, impurities are adsorbed in an adsorber at higher partial pressure and then, desorbed
at lower partial pressure. By using this process, hydrogen is recovered at high pressure and less
impurities, because very little hydrogen is adsorbed relative to methane and other light hydrocarbons.
Commercial scale PSA units normally use between 4 and 12 adsorbers. More adsorbers are used for
higher hydrogen recovery or increasing capacity. The driving force for separation in this process is the
difference in impurity partial pressure between the feed and tail gas. Two major advantages of this
pro ess are it’s a ilit to produ e a high pressure a d high purit e ess of 99 ol% and frequently
99.99 vol% hydrogen stream [1-3]. In this process amount of hydrogen recovery is moderate (65-90%
depending on the tail gas pressure) because a part of produced hydrogen is consumed for regeneration
the beds. PSA system can be significant and the operating pressure of a PSA unit should be optimized.
PSA systems are insensitive to feed change, robust, efficient and normally used in many refineries across
the world.
2.2. Cryogenic process
The cryogenic process based on partial condensation, removes the hydrocarbon impurities from the
hydrogen stream. Cryogenic units are based on the difference in volatility (boiling temperature) of the
feed components. Hydrogen has a high relative volatility compared with methane and other light
hydrocarbons. In this process, the required amount of feed impurities is condensed by cooling the feed
stream against warming the product and tail gas streams in multi-pass heat exchanger. The refrigeration
required for the process is obtained by Joule-Thomson refrigeration derived from throttling the
condensed liquid hydrocarbons. This process is typically applied for separation of hydrogen-
hydrocarbon. If the feed contains water and other components that could freeze in the system, rather
than entering to cryogenic unit should be preheated.
Thus the cryogenic unit, typically separate the feed into three products, a high purity hydrogen stream,
methane-rich stream, and C2+ hydrocarbons product. Additional products, such as ethane-propane and
LPG, can be produced using additional separators. When the feed pressure is low, the feed hydrogen
content is less than 40% and there are higher concentrations of heavier hydrocarbons which can be
easily condensed, cryogenic process can be the best process for hydrogen purification[1-3]. Cryogenic
process is cost intensive and in processing varying feed composition has less flexibility and sometimes
requires supplemental refrigeration and is considered less reliable than PSA or a membrane process.
Due to disadvantages of cryogenic process for hydrogen purification, this technology is very rarely used
in refineries.
2.3. Membrane process
Membrane systems are based on the difference between in permeation rates between hydrogen and
impurities across a gas-permeable polymer membrane. Permeation involves three sequential
mechanisms: the component of gas phase must first adsorbed/dissolve into the membrane (higher
pressure), diffuses through in membrane matrix and desorbed it to the permeate side (lower pressure).
Different components have different solubility and permeation rates. Solubility depends on the chemical
composition of the membrane and diffusion on the structure of the membrane. Components with
higher permeability, such as hydrogen, dissolve in to the polymer membrane on the high pressure side
and diffuse to the low pressure side and components with lower permeability, are retained on the high
pressure side because of the depletion of components with high permeability. High permeation rates
are due to high solubilities, high diffusivities, or both. The driving force is the difference in partial
pressure, with the highest driving force giving the highest recovery.
The polymeric membranes used for separation are consists of cellulose acetate, polyacetate,
polysulfonate, polyamide and polyimide. Membrane units can recover hydrogen at moderate purity (9097%) and moderate recovery (85-95%) [3].
Polymer membrane technology appears as an attractive alternative for the separation of gases in
industrial processes, compared with conventional separation technologies. Membranes exhibit
simplicity of operation and maintenance, small size, efficient and reliable performance. These appealing
features enabled membrane technology to penetrate wide variety applications. Today, membrane
technology is successfully applied at industrial level for hydrogen recovery from ammonia purge
streams, syngas ratio adjustment and hydrogen recovery in refineries. These three applications involve
easy task separations since the extraordinary small molecular size of H2 makes it extremely permeable
and easily collected as permeant compared to the other slower permeating components: N2, CO and
CH4. Membranes have no moving parts and are reliable compared with other available processes [4-5].
3. Technology selection for hydrogen recovery from refinery off-gases
The hydrogen content and operating pressure of the refinery off-gasses have a large influence on both
the process recovery selection and the capital investment for the recovery unit. Table 1, depicted
presents hydrogen content (vol%) and pressure of various hydrogen-rich gases that are obtained in
petroleum processing.
Table 1: Hydrogen content (vol%) and pressure of various hydrogen-rich gases in refinery
Process
Catalytic reforming
Catalytic cracking (off-gas)
Hydro-cracking (purge)
Hydro-treating (purge)
Hydrogen content (%vol)
40-85
10-20
40-80
25-40
Initial pressure (MPa)
2.5
5.5
1.2-2.85
4.0
In cryogenic, requires the feed with relatively lower hydrogen purity (typically 20-40 vol%) and hydrogen
pressure loss is much less than membrane systems. In cryogenic distillation units, recovering of byproducts, such as ethane and methane is possible. The cryogenic process is thermodynamically the most
effi ie t h droge purifi atio te h olog , ut PSA pro ess, despite it’s lo er h droge re o er , is the
most commonly used in petroleum refinery due to better performance, ease of operational and
economical factors. Cryogenic distillation is considered highly energy intensive and expensive separation
method compared to other process [3-6]. A broad comparative evaluation of technologies was done and
the same is enclosed in Table-2.
Table 2 - Process and operational consideration for hydrogen purification technologies
Parameters
Cryogenic
Membrane
PSA
H2 feed (min), %
15
15
50
Feed pressure, psig
200-1,200
200-2000
150-1000
H2 Purity, %
97+
95+
99.9
H2 recovery, %
Upto 98
Upto 97
Upto 90
CO+CO2 removal
No
No (only CO)
Yes
H2 product pressure
App. Feed
Much less than feed
App. feed
Feed pretreatment
Yes
No
No
Flexibility
Average
High
High
Reliability
Average
High
High
By-product recovery
Possible
Possible
No
Ease of expansion
Low
High
Average
Energy requirement
High
Low
Moderate
Membrane separation is an emerging technique due to certain advantages such as modular
nature (thus easy scale up), lower foot-prints, simplicity in operation, better separation at ambient
conditions etc. Mostly, polymeric membranes are used in industrial scale gas separation due to their
lower cost, ease of preparation, acceptable separation performance, etc. Polymeric membrane based
separation could be better option for H2 recovery from purge gas or fuel gas streams in refinery due to
its high selectivity of H2 recovery over other hydrocarbon mixture. [7-8].
4. Indigenous Membrane Development
Present status in India
Various polymeric membranes are known for the gas separation. Although commercial membranes are
available for this purpose, there are various issues associated with them such as (i) high cost (ii)
availability of only large industrial scale modules / systems (iii) feeble technical support from suppliers,
etc. These issues make it difficult to employ membranes in Indian scenario with our own technocommercial challenges.
5. Scope of the work
Many polymers are reported in the literature for H2 separation [7-9]. Membrane material selection
criteria for a required separation application are dictated by designing (manufacturing) and operating
parameters of membranes. All these aspects finally decide separation performance, long term
membrane stability, cost feasibility, etc. These factors needs to be considered while selecting a
membrane material for present application of H2 recovery from the mixture of gaseous hydrocarbons.
Recent, study shows there are some Indian institutes like CSIR-NCL and CSIR-CGCRI in India who has
developed the technology of making membranes in different shapes and sizes using readily available
materials and the cost has drastically come down (~1/10th cost reduced) compared to imported units.
Present paper, deals with the performance study of selected polymeric membrane based on refinery
synthetic feed and optimization of process parameter for better recovery and purity of hydrogen gas.
6. Experimental procedure
Membrane bench scale unit is equipped with a polymeric membranes supplied by module supplied by
one of the membrane manufacturers who used to make polymeric membranes using advanced spinning
technique for oxygen/ nitrogen separation membranes for medical applications. This manufacturing
technique was adapted for H2 separation from hydrocarbons. The module contains combination of
different hollow fibers membranes with effective surface area varies from 0.1-1 m2. The design
specification of membrane modules is given in table-3. The membrane module can withstand pressures
and temperatures of 10 bars and 100 °C, respectively, while it can handle gas stream flows of up to 0.55 Nlit/min. Fig. 1, shows the flow scheme of membrane experimental set up. The membrane unit
contains all necessary ancillary instrumentation for pressure, temperature, flow rate measurement and
control. The feed pressure is maintained by a pressure regulator installed in the source of gas stream.
Both permeate and residue flow rates are measured after expansion to atmospheric pressure using wet
gas meter. The feed gas enters the shell side at high pressure and flows inside the fibers in a counter
current mode to the permeate flow. The composition of each gas stream is analyzed by Refinery gas
analyzer (make Agilent 7890A) Gas Chromatograph equipment.
Fig.1: Schematic of lab scale bench scale unit for membrane testing
7. Result and discussion
The composition of multi-gases feed being used in membrane performance is given below in Table 4.
The selection of feed is based on the feed composition is available in various plant in refinery.
Table-3: Specification of membrane modules used in the screening of membranes
Composition
MPH-2
Effective Membrane Area (m2)
Diameter (inch)
Type
Selectivity (H2/CH4)
0.1-1
½’ to
Hollow fiber
40
Table 4: Feed composition of different feed used in membrane screening
Composition
Vol%
Composition
Vol%
Hydrogen
Propane
Ethylene
I-butane
40-80
2-5
1-5
1-5
N-butane
Ethane
Methane
Propylene
2-5
1-20
1-25
1-3
7.1. Selection of membrane module
Table 1, summarises some potential sources for hydrogen recovery in a typical oil refinery. In all cases,
the concentration of hydrogen varies from ~20 to 80%, the stream temperatures are relatively low (30–
40oC) and the pressure varies from 10 to 40 bar [3]. Undoubtedly, these values of major process
variables are all compatible with polymeric membrane separation technology.
The gaseous mixture of was fed to the hollow fiber membrane unit at constant pressure, operated at
different retentate flow rate (400-1200 ml/min i.e., different fraction of stage cut). The feed flow rate
are varied from ~500-5000 ml/min and gaseous concentrations in permeate and residue streams were
measured as a function of stage cut. In all cases, flow rates and concentrations were measured with an
accuracy of +/- 0.5 and +/- 1.5%, respectively.
Fig. 2A- 3B, shows the variation of H2 recovery and H2 purity with flow rate (permeate and retentate) at
constant pressure. Figure shows that hydrogen is enriched in the permeate stream from 80 to 95vol%
and 60 to 90% at various retentate flow rates at same operating pressure conditions. For feed-2,
maximum hydrogen recovery (<90%) and purity (<95vol%) with maximum permeate flow rate was
observed at moderate flow rate of retentate. This condition is optimized based on the higher permeate
flow rate, higher H2 purity and moderate recovery of hydrogen for further lab/scale up operations in
refinery.
4000
1600
95
1000
800
90
85
600
400
80
Flow rate (ml/min)
1200
98
3500
Recovery(%)
Flow rate (ml/min)
1400
100
4500
100
2500
94
2000
92
1500
500
0
0
Fig.2A Enrichment (H2 purity) and recovery of
hydrogen in permeate stream with varying
retentate flow using feed-1.
90
1000
200
75
91
89
87
84
74
Purity(%) , 4 [email protected]
Permeate
Retentate
Recovery(%)
96
3000
Recovery (%)
1800
88
86
96 95.3 94 93 90.5 89 86
Purity(%) @4 [email protected]
Permeate
Retentate
Recovery(%)
Fig.2B Enrichment (H2 purity) and recovery of
hydrogen in permeate stream with varying
retentate flow using feed-2.
8. Conclusion
The majority of gas separations encountered in refinery/industrial level refer to multi-component
mixtures. Among them the separation of hydrogen from light hydrocarbons in refineries is an issue of
major importance in terms of economic and environmental concern. In this regards, polymer
membranes can be viewed as an attractive alternative option over a PSA or cryogenic processes due to
modular nature (thus easy scale up), lower foot-prints, simplicity in operation, better separation at
ambient conditions etc. Mostly, polymeric membranes are used in industrial scale gas separation due to
lower cost and ease of operation, acceptable separation performance and robust technology. In that
direction, process optimization and selection of right kind of polymeric membrane materials can
powerful tool for process integration in refineries. Although, membrane material selection criteria for a
required separation application is dictated by nature of feed gas, anticipated selectivity, plasticization
aspect, long term thermal/mechanical stability, etc. These parameters are important along with
availability of polymers on industrial scale. In this study, results show that even for a one-stage
membrane unit, high hydrogen permeate purity (<95 vol%) and significant total recovery (<90%) can be
achieved for moderate to high permeate flow (to reduce the capex cost) and stage cuts. The residue
stream, rich in hydrocarbons and containing left over hydrogen can be further separated in a second
stage, or used as a fuel gas.
References
1. R.D. Noble, S.A. Stern, Membrane Separations Technology, Elsevier, Amsterdam, 1995
2. S.P. Kaldis, G.C. Kapa taidakis, G.P. Sakellaropoulos, Si ulatio of ulti o po e t gas
separation in a hollow fiber membrane by orthogonal collocation — hydrogen recovery from
refi er gases , Jour al of Me ra e S ie e
–71.
3. Zahra Rabiei, Hydrogen management in refineries, Petroleum & Coal 54(4) 357-368, 2012
4. A.F. Ismail, K.Ch. Khulbe,Takeshi Matsuura, Gas Separation Membrane- polymeric and inorganic,
Springer (book), ISBN 978-3-319-01094-6.
5. Lu Shaoa, B. T. Lowa, Tai-Shung Chunga, A.R. Gree erg, Pol eri
e ra es for the
h droge e o o : Co te porar approa hes a d prospe ts for the future , Jour al of
Membrane Science 327 (2009) 18–31
6. R.W. Baker, Membrane Technology and Applications, McGraw-Hill, New York, 2000
7. Richard W. Baker, Future Directions of Membrane Gas Separation Technology, Ind. Eng. Chem.
Res. 2002, 41, 1393-1411.
8. Zachary P. Smith, Rajkiran R. Tiwari, Thomas M. Murphy, David F. Sanders, Kristofer L. Gleason,
Donald R. Paul, Benny D. Freeman H droge sorptio i pol ers for e ra e appli atio s
Polymer 54(2013) 3026-3037.
9. Lonsdale HK. The growth of membrane technology. Journal of Membrane Science 1982;10(23):81-181
Acknowledgement
The authors express sincere thanks to the BPCL management for providing necessary support to conduct
the research work at Corporate R&D Centre, Bharat Petroleum Corporation Ltd., Greater Noida.
Opportunity Crudes (High Calcium) – Processing challenges
(By: Nagashyam Appalla (Reliance refinery division))
Abstract :
Opportunity crudes are available at a discounted price either due to problems associated with
processing such crudes or due to lack of any previous experience in processing them. Inclusion
of such crudes in crude basket reduces the price of crude oil, however extreme care needs to be
taken while processing them to ensure minimum impact on Crude & downstream secondary
processing units. A comprehensive processing strategy needs to be devised while processing
such crudes. Crudes containing calcium in the form of calcium napthenates are opportunity
crudes. They are available at discount compared to the conventional crude oil. Typical buying
margins on these crudes could be ~ $ 1-2 per barrel. Calcium napthenates are natural emulsion
stabilizers that could stabilize the oil water emulsion in crude desalters & lead to process upsets.
Most of the calcium in desalted crude lands up in short residue leaving vacuum tower which could
have a detrimental impact heater performance in downstream Coker unit. For refineries feeding
short residue directly to RFCC the increased calcium in feed acts as catalyst poison leading to
deactivation of catalyst. Three options are available for processing such high calcium containing
crudes viz. i) Removing the calcium from crude before processing (Not well established
commercially) ii) Restricting the % of such crude to very low % in blend (Lowers margin) & iii)
Comprehensive chemical treatment program to remove calcium in desalter (Commercially
proven). The chemical treatment program to remove calcium from crude oil comprises of
acidification of desalter wash water using weak organic acid that would react with the calcium
napthenate to form calcium salt of acid which is soluble in water. This moves the calcium from
hydrocarbon phase to water phase in desalter. The required quantity of acid depends on the target
calcium removal. Calcium that moves to desalter brine could pose issues in downstream brine
treating unit or effluent treatment plant processing the desalter brine. The problem becomes more
profound for refiners with zero discharge/complete recycle of water.
This paper covers strategy & best practices that can be adopted to minimize the adverse impact
& maximize margins while processing crudes containing calcium napthenate which includes
criteria for selecting acid, hardware requirements, crude blending strategy, monitoring parameters
& key performance indicators. It also covers the expected problems in Crude & Downstream units
which processing such crudes and ways to manage and mitigate them.
What are Opportunity crudes? :
Opportunity crudes are crudes which are available at discounted price either to problems
associated with handling & processing them or due to lack of any previous processing history.
Though inclusion of these crudes in the crude basket reduces the cost of feed and improves GRM
(Gross Refining Margin), a conscious call needs to be taken while adding these crudes in basket.
Opportunity crudes come with challenges in handling & processing them without impacting the
integrity & product qualities. Some of the opportunity crude & the problems associated with them
are listed in Table-1 below.
Property
TAN
Metals
API
Calcium
(As Calcium Napthenate)
Viscosity
Filterable solids
H2S
Mercury
Issues in processing
High - Napthenic acid corrosion attack.
- High conductivity, poor desalting.
- Catalyst deactivation.
- Poor desalting with existing hardware.
Low
- High oil under carryover in brine water.
- Emulsion stabilizer.
- Impact on downstream Coker heater.
High
- Fouling/Choking of desalter brine handing system.
- Catalyst poison.
- Poor desalting.
High - Poor mixing of oil & water.
- Need special handling facility.
High
High
- High Oil under carryover in brine.
- Fouling of exchangers.
High - Issues in handling water draining of tanks.
High
- Impact of biological growth in ETP.
- Catalyst poison in hydrotreater.
Table-1: Opportunity crudes & issues in processing them.
Why Crudes containing Calcium are Opportunity crudes?
Some crudes contain high amount of organic calcium (~ 150 – 300 ppmw) in the form of Calcium
Naphthenate. Most of these crudes are typically high TAN crudes. Calcium Napthenates are
emulsion stablizers (surfactants) which stabilize the emulsion of oil & water in desalter leading to
poor separation. They have low solubility in oil & water and form napthenate soap layer at
interface. Conductivity of these crude is high which leads to high current drawn in the desalter
electric grids & voltage decay. This in turn leads to carry over of higher BS&W (Sediments &
water) at the outlet of desalter & high oil under carryover in brine. Most of the calcium in crude
lands up in the short residue leaving the vacuum tower which has detrimental impact on the
downstream units. For refineries feeding the short residue to a RFCC the increased calcium acts
as a catalyst poison leading to de-activation of catalyst. It also impacts the quality of fuel oil due
to increased calcium leading to slagging in furnace burners. For refineries that feed the short
residue to a Coker, the impact of increased Calcium loading is seen on the quality of coke. Also
increased calcium is found to aggravate the tendency of coking/fouling in coker fired heater tubes
leading to increased skin temperatures.
Due to above challenges associated with processing, high calcium crudes qualify as opportunity
crudes. Typical buying margins on these crudes is ~ 1–2 $/bbl and hence very lucrative for
inclusion in the crude basket.
What are the options to process Calcium containing crudes?
Three options available for processing these crudes is listed in Table-2 below.
Option-1
Option-2
Option-3
Option
Remove the calcium from crude
before processing using some
technology (eg. Ion-Exchange).
Restrict the crudes to a very low % in
the blend to avoid any impact
Go for a chemical treatment program
that enables removal of calcium in
desalter along with brine.
Remarks
Such methods have not been established
& proven commercially.
This would increase processing time &
may reduce margins.
This is commercially proven method as
on date
Table-2: Options to process Calcium crudes.
What is chemical treatment for removing Calcium in desalter ?
Desalter acidification program has been proved to be successful in removing metals like Ca, Fe
from crude in desalter. The acidification program involves injecting an acid in wash water system
of desalter to lower the pH of desalting water. Acidification also helps in resolving the difficult oil
water emulsion and trapping the tramp amines that come along with crude. Typically a weak
organic acid is used for acidification purpose. The chemical reaction involved in the removal of
Calcium is as under.
Ca-Napthenate
(Oil soluble)
+
Organic acid 
(Water soluble)
Calcium – Organic acid
(Water soluble)
+
Naphthenic acid
(Oil soluble)
The organic acid reacts with the calcium napthenate and forms a salt of calcium that is soluble in
aqueous phase liberating the Naphthenic acid molecule which goes along with the hydrocarbon.
What is required for carrying out chemical treatment ?
Selection of acid for treatment.
Acid that are used for acidification of wash water in desalters can be either In-Organic/mineral
acids (H2SO4) or Organic acids.
In-organic acid:
In-Organic acids like H2SO4 (Sulfuric acid) are cheap and easily available. However use of H2SO4
may lead to carry over of strong acid species to distill in atmospheric tower overhead which may
aggravate the corrosion in overhead system. H2SO4 has tendency to form SO4-2 salts in brine
water which can deposit in the brine system leading to severe fouling. Also H2SO4 being a strong
acid, there is a high potential of localized corrosion at the injection point which may warrant use
of higher metallurgy like Haste alloy at the injection point. Industry wide the use of H2SO4 for wash
water acidification in desalters is very low due to various issues experienced with it’s use.
Organic acid:
Commercially available weak organic acids like acetic acid, glycolic acid, citric acid etc. are widely
used for acidification of wash water. Use of organic acids for pH correction in desalters for
improving the desalter performance is well known. Care needs to be taken while selecting the
acid for treatment program based on the partioning of acids in hydrocarbon and the solubiliy of
salts of acid in water.
Some acids tend to partition more in the hydrocarbon phase leading to carry over of acid species
to atmospheric tower overhead which increases the requirement of neutralizer in the overhead
system. Some acids have very poor solubility for calcium salt formed with acid. This may lead to
precipitation of calcium salts either in desalter or downstream brine circuit at lower temperature.
Chemical vendors treating desalters offer proprietary formulation of organic acid which may also
contain corrosion inhibitor to prevent low pH corrosion at injection point and calcium dispersant
to prevent deposition of calcium salts in the brine system.
A proper evaluation should be done while selecting the acidification program based on the system
requirement. COD (Chemical oxygen demand) of the acid should be considered during evaluation
to assess the impact on ETP (Effluent treatment plant).
Hardware requirement for acid injection:
The injection of acid for acidification is normally done in desalter wash water system. A dedicated
dosing skid with metering pumps would be required to inject the acid in wash water going to both
1st and 2nd stage desalters. The piping for the acid injection should be SS to avoid any corrosion.
It is recommended to have a higher metallurgy lining downstream of injection point for some length
to avoid localized low pH corrosion. Injection should be done using injection quills which should
be either SS316 or higher metallurgy.
Normally the chemical dosage is based on the calcium content in the raw crude. Fine adjustment
of the chemical is done based on the actual removal rate of calcium, oil in brine and BSW in
desalted crude. The metering pumps shoud be capable of handling the acid requirement based
on the dosage rate required for range of calcium content in feed.
What is the strategy for processing ?
PED (Portable Electric Dehydrator) testing of blend:
Lab scale study of oil/water separation, rag layer formation and sharpness of interface can be
done with demulsifier and acidified wash water before actual processing. This would help in
identifying the best demulsifier suited for processing the blend.
Blending:
Processing should be first started with a very low % in blend to check the impact on system. The
% of calcium crude shoud be increased in blend gradually after evaluating the performance at
each step change and collecting various data points. Care should be taken to ensure that crudes
containing tramp amines (that can lead to high pH) are not blended while processing these crudes
as they would increase the requirement of acid.
Monitoring:
A more robust monitoring is required while processing high Calcium crudes. This includes
rigorous monitoring of various parameters affecting the performance of desalter & downstream
units, carrying out additional analysis of samples of crude, brine & product streams, carrying out
more frequent field checks on desalter performance. A typical list of checks to be carried out
during processing of high Calcium crudes is given in table below.
Type Check
Analytical
Online
monitor
Parameter
BSW in desalted
crude
Salt in desalted
crude
Calcium in crude
blend
Calcium in
desalted crude
Calcium in VR and
Gas oil
UOM
% vol
ptb
ppmw
ppmw
ppmw
Grid
voltage/Current
Volts/Amps
Interface level
%
Corrosion rate of
brine header
mpy
pH of brine
Agar probe in
desalted crude
%
Overhead boot
water
pH
Agar probe in
brine
%
Remarks
With Ca in crude blend the desalter performance
may get impacted due to formation of a more
stable emulsion. This may lead to increased
BSW & Salt in desalted crude.
Calcium in inlet and outlet of desalter should be
monitored periodically to estimate the removal
efficiency across the desalter.
Calcium in VR and Gasoil should be analyzed
periodically to check the increased Ca loading.
Increase in conductivity leading to higher current
drawn in the grids.
Voltage decay leading to poor desalter
performance.
Continuous trending of voltage & current may be
done and record the changes with increased
calcium in blend. The trends may give an early
indication for potential upset and taking
corrective actions.
Desalter grids to be kept on lowest voltage tap.
The interface level should be maintained high
enough to provide sufficient residence time to
brine. This helps to reduce the oil under
carryover in brine.
Trending online probe would give an indication
of corroion rate in brine header due to injection
of acid.
Trending online pH indicator would give
guidance for adjusting the brine pH.
Trending may provide early indication of any
water shots to take corrective action.
Trending online pH indicator would give
indication of portioning of acid to oil & carryover
to atmospheric tower and leading to higher
dosage of neutralizer.
Trending may provide any abnormal ingress of
oil in water.
Trycock samples
visual
Brine appearance
visual
Acid injection rate
lph
Physical
check
Try cock samples should be checked regularly
to detect any rag layer build up. In case of any
such indication shock dosage of demulsfier may
help to break the rag layer. Alternately
skimming of interface layer may be done where
facility is provided.
Frequent check of brine should be carried out to
detect any abnormal carryover of oil in water.
Injection of acidified wash water from desalter
to preheat train may be reviewed and stopped if
required.
Dosage rate should be checked periodically to
ensure no suspension of dosage which may
cause upset of desalter.
Table-3: Additional checks during processing Calcium crudes.
Key Performance Indicators (KPIs):
Typical performance metrics that can be agreed upon with desalter treatment vendor during
processing of crudes containing calcium napthenate are as under.
BSW in desalted crude
Salt in desalted crude
Oil in water
UOM
Standard desalter KPIs
% vol
ptb
ppmw
Remarks
Additional KPIs while processing Ca crudes
Calcium removal efficiency across desalter
%
CRA consumption
ppmw/ppm of Ca
Acidic corrosion in wash water and
brine system
Corrosion rate in brine line
mpy
Increased pressure drop of brine header
kg/cm2
Fouling of brine system
Table-4: Typical KPIs during processing Calcium containg crudes.
What is the impact on downstream units?
Though these crudes can be processed using a chemical treatment programme and a robust
monitoring system, some problems associated with processing them cannot be eliminated
Brine system:
The calcium gets removed from the crude in the form of calcium salt of organic acid which
migrates to the water phase. It is very important that the calcium salt is soluble in brine and doesn’t
precipitate out. The solubility of the salt in brine is sensitive to pH of brine. Higher pH favors
precipitation of salts which may deposit down in the brine header, brine wash water exchangers,
run down aerial coolers. This could increase the back pressure of the brine header and increase
the opening of brine control valves. Some vendors recommend use of Calcium dispersant to
prevent the salts from settling down in the system. The impact could be increased cleaning of
the brine-wash water exchangers and aerial coolers and requirement to clean the brine header.
Addition of caustic in brine header or brine stripping should be avoided while processing these
crudes to prevent precipitation of calcium salts. Back pressure of brine system should be
monitored for any precipitation & fouling.
ETP:
It has been observed that the organic acid that is injected in desalters binds NH3 to some extent
which is thermally stable and doesn’t get stripped off by steam stripping. This leads to increase
load of NH3 in ETP influent while processing these crudes. The impact is increased load on the
biological system in ETP and increased COD of treated water. For refiners that have a RO
(Reverse Osmosis) system the increased calcium content in water increases the rate of
regeneration of pre-treatment ion exchange beds and may cause fouling of RO membranes.
RIL processing experience:
RIL has been successfully processing crudes containing calcium napthenates using wash water
acidification program since 2012 without any significant impact on any units. Below is the trend of
quantity of crudes containing calcium napthenates processed over years.
Summary:
-
Crudes containing high amount of Calcium in form of Calcium Napthenates are opportunity
crudes and can fetch a typical processing margin of ~ 1-2 $/bbl.
Acidification of wash water in crude desalters is proven way to shift the calcium from crude
oil (hydrocarbon phase) to brine water (aqueous phase).
A good chemical treatment programme & robust monitoring process is required while
processing these crudes to limit issues in crude and downstream units.
Fouling of brine system and impact of downstream biological system remains a challenge
while processing them.
These crudes can be considered in crude basket if the margin in processing them is
attractive even after deducting the cost of additional chemical treatment & maintenance.
Reference:
Technical paper: Desalter acidification additives and their potential impacts on crude units –
NACE-Corrosion 2008 Conference & Expo.
Crude to Petrochemicals – A New Paradigm
Sanjay Rajora, Jt. General Manager;
Ms Manasi Patel, Senior Manager;
Hillol Das, Vice President;
Essar Oil Ltd., Mumbai
Introduction
The petrochemical industry is a key enabler of modern living. Ethylene, Propylene and
Aromatics are major building blocks to derive various petrochemical derivatives. A conventional
approach for building petrochemical facilities is either with integration with a refinery or with a
specific feed for petrochemicals. Rarely a refinery is designed with an approach to maximize
petrochemicals. It has always been a balance between fuels and petrochemicals or rather
inclination towards fuels if refinery is built first. It is observed that refining margins are volatile
and it poses serious challenges to refiners looking at the ongoing developments in the global
business.
Right feedstock to get the final desirable products from a petrochemical facility gives a
competitive advantage as it covers a major part of the cost. A feedstock which is cheaper and
easily available provides strength to the project economics. A study was taken up to analyze a
petrochemical facility built on heavy crude as a primary feedstock and targeted to maximize
petrochemicals with minimizing the fuels production. Subsequent sections present the approach
adopted in this paper and options studied for designing the petrochemical facilities.
Considerations
We have taken an approach to have a blend of traditional and advanced technologies to
achieve the objective of maximizing the petrochemical molecules through a refining facility.
Following considerations are taken for the study:
-
15 MMTPA heavy crude mix of about 24-25 API as main feedstock to the complex
-
Conventional primary and secondary treatment for light and middle distillates
-
Using advance technologies for VGO and VR conversion
-
Maximizing the feed to cracker by using all the lighter molecules
-
Minimizing the fuels production
-
Production of petrochemical building blocks like Ethylene, Propylene, Benzene, and
Paraxylene
To facilitate the study, a Linear Programming based model was used to optimize the
configuration and a financial model was used to analyze the attractiveness of the selected
schemes. Two selected schemes are being considered here for the discussion.
Scheme -1
The basis of the scheme is to have a 15 MMTPA crude refining facility followed by necessary
treatment facilities for light and middle distillates in conventional way. Post hydro-treatment, light
molecules like LPG and Naphtha are sent for cracking in Steam Cracker Complex (SCC). VGO
is hydrotreated before being fed to high propylene fluidized catalytic cracking unit (HPFCC)
thereby generating maximum amount of light olefins and Naphtha rich in Aromatics. Light Cycle
Oil from FCC is also cracked to get feed for Aromatics.
For VR conversion, Solvent Deasphalting (SDA) and Slurry Hydrocracking (SHK) are
considered. Deasphalted oil from SDA is treated to make suitable feed for FCC. SHK is run on
Pitch from SDA and neat products from SHK are sent for further conversion.
The scheme covers three major sources for petrochemical molecules generation. The first one
is High Propylene Fluidized Catalytic Cracking Unit, second is Aromatics complex (ARX) and
the third is Steam Cracker Complex. All the C3s/C4s and Light Naphtha are sent to Cracker
Complex. Heavier part of Naphtha is reformed for conversion to Aromatics. The facility is
considered to be built in such a way that all LPG and Gasoline range molecules are being
converted into petrochemicals. Middle Distillates are the major fuel products produced in the
process. The major products from the scheme are mentioned in the table below:
Major Products
Million Ton per Annum
Ethylene
1.6
Propylene
1.8
Paraxylene
1.5
Benzene
0.5
Distillate (Diesel + ATF)
5.8
It may be observed that significant amount of petrochemicals are produced from the scheme.
The Petrochemicals produced are about 36% to the crude intake which is quite high compared
to conventional integrated petrochemical complex.
Scheme – 2
The only difference between scheme-1 and scheme-2 is not having a SDA unit in scheme -2. It
results in little lower feed for FCC unit and hence slightly lower production of petrochemical
molecules. The table below presents the product slate from scheme -2.
Major Products
Million Ton per Annum
Ethylene
1.4
Propylene
1.4
Paraxylene
1.4
Benzene
0.5
Distillate (Diesel + ATF)
6.7
The Petrochemicals to crude are 31% in this scheme. A common schematic is shown below for
both the schemes with the red highlighted portion (SDA loop) applicable to scheme-1 only.
Analysis & Discussion
It is very interesting to note the significant petrochemical building blocks production through both
the schemes discussed above. The analysis has shown the P/E (Propylene to Ethylene) ratio of
1.0 or above in the options studied vis-a-vis about 0.1 to 0.6 in a conventional steam cracker
based on various feedstock. This is important from the fact that propylene is growing at a faster
rate than the ethylene. Also, as crackers are switching to lighter feed, the propylene production
from crackers is falling. In such a scenario, more propylene production from a petrochemical
complex is a welcome move.
Investment for such a facility may require about $6.0 to $6.5 billion for ISBL part as per in-house
data. In the analysis, the scheme -1 is found to be costlier by about $700 million on plant and
machinery cost basis. This is comparable with any large petrochemical complex of the size
discussed above. The key point here is to get more petrochemicals with cheaper feedstock.
However, the facility can be designed for lower capacities also as per the investment appetite of
the investors. Product slate optimization can be done with specific technological input taken
from the technology providers.
From profitability point of view, historical price sets for a certain period are considered to
calculate the margins for the schemes under study. Both the schemes are found very attractive
and Gross Margin was calculated to be higher by about $6-$10 per barrel in comparison with
the fuels only schemes.
Conclusion:
The above schemes suggest an alternative way to produce a good mix of petrochemical
molecules from crude. The yields of petrochemical building blocks are significantly high and
provide a good feedstock for building a large scale petrochemical complex. The schemes
discussed are found to be economically attractive and may be taken up for further studies while
putting up petrochemical facilities.
Disclaimer: The scope of this publication is strictly for knowledge sharing purposes and not necessarily to provide
any recommendation to the audience/readers. Any statement, opinion, or observation made in this presentation are
those of the presenter only and do not represent the intent and business plan of Essar Oil Ltd (EOL) nor the figures
represented here are verified by any statutory agency/institution.
HP Gelators – Organic Soft Materials – A Potential Solution
for Oil Spillage Remediation and Recovery
Chinthalapati Siva Kesava Raju, Bhaskar Pramanik, Kandanelli Ramesh, Mangala Ramkumar,
Raman Ravishankar, Peddy Venkat Chalapathi Rao and Gandham Sriganesh
Hindustan Petroleum Green R&D Center (HPGRDC), KIADB Industrial Estate, Tarabahalli,
Hoskote Taluk, Bangalore, 560067, Karnataka, India, Email: [email protected]
Keywords: organogelator, dipeptide, crude oil, oil spill
Abstract
Dipeptide based four compounds 1-4 were synthesized using Valine and/or Methionine amino
acid along with long chain fatty acid Palmitic acid. The synthesized compounds were subjected
to phase selective gelation of oil phase in a biphasic mixture of oil and water. A wide range of
organic solvents, mineral oils including crude oil of varying APIs were effectively transformed
to gel state using compound 1-4 where the uptake capacity of the oil phase was found to be in the
range of 40 to 125 times of the gelator weight. Conversion of the dipeptide compounds to their
corresponding sodium salts was found to shift the gelation property towards (aquo) hydrophilic
medium. A broad range of applicability, effectiveness for a wide spectrum of oils and crude oils
dictates these compounds as one of the potential candidates for the oil spillage recovery.
Introduction
The oil spillage has become a common problem from well to wheel i.e., from drilling of oil to
different stages of transportation.1 The world has witnessed several oil spillage incidents and one
of the major incidents in recent time is the 5 billion barrels of crude oil release in the Gulf of
Mexico and also the ship accident in Chennai, the south east coast of India leading to huge oil
spillage spreading across the east coast of India. The adverse effect of oil spillage is associated
with environmental pollution and is life threatening to marine ecosystem.2 Development of new
methods for containing the oil spills is the need in terms of valuable economy and environmental
view point. Conventional methods for oil spill recovery include combustion, mechanical
treatment using oil sorbent materials, chemical treatment by dispersants, and bioremediation with
microorganisms.3 All of these processes are associated with certain drawbacks such as. poor
recovery, uneconomical, release of toxic residues, time consuming and so forth. Recently, low
molecular weight organogelators (LMOGs) have gained immense interest as an alternative to
mitigate the oil spills.4 Small molecule gelators are recognized for their ability to immobilize
solvent molecules into their self-assembled fibrillar networks.5 The nano-scale or micro-scale
network structures are formed by non-covalent interactions like hydrogen bonding, π-π stacking,
van der Waal interactions, dipole-dipole interactions, and charge-transfer interactions.6 Gels are
solid-like network in liquid-like medium which can display the rheological properties of solidphase materials, high surface area for liquid phase and rapid internal diffusion kinetics for self
healing.7 These distinct properties of gel dictates vast range of applications in catalysis,
pharmaceuticals, bio-technology, light harvesting, lubricants and food industry.8 However for oil
phase gelation and recovery purposes, the phase selective gelator must fulfill some requirements
such as: (a) Cost effective and ease of synthesis (b) selective and efficient gelation of oil phase,
(c) rapid gelation at room temperature, (d) stable gel with respect to temperature and external
force, (e) easy oil recovery from the gel phase, and (f) reusable. Thorough understanding of the
supramolecular gelation mechanism led to rationally designed small molecule gelators with
desired functionality to serve a specific task of gelating the oil phase selectively. In this regard,
amino acid-based LMOGs are frequently used where proper balance between hydrophilic and
hydrophobic functionalities in the molecule dictated the phase selective gelation behaviour.
Amino acid based peptide molecules are known to exhibit gelation of hydrophilic solvents as
well as hydrophobic solvents/oils depending on their structural features.9 Long hydrophobic
alkyl chain attached to the peptide molecule can direct the molecule to be oil selective. Thus,
proper tuning of the peptide molecule to be capable of selectively rigidifying the crude oil and
other refinery distillates from a biphasic mixture of the hydrocarbons with sea water can find
realistic application in oil spill/recovery.
Herein, four dipeptide-based amphiphilic LMOGs 1-4 (Scheme 1) are reported for selective
gelation of crude oils of varying API other refinery products (SRN, CRN and Diesel) as well as
organic solvents from a biphasic mixture of oil/water.
Scheme 1 Structure of the dipeptide gelators
Results and Discussions
Amino acid based amphiphiles were synthesized by conventional solution phase methodology as
reported previously.10 These compounds differ with each other based on amino acid (Methionine
and Valine) functionality.
Figure 1 Photographs of gels obtained with different oils (C1 and C5 are crude oils according to
Table 2) in a biphasic mixture with water
The gelation ability of the compounds were examined in different solvents (n-paraffins, alkene,
aromatics), refinery fractions like SRN, CRN and diesel, crude and vegetable oils by ‘stable to
inversion of glass vial’. Gelation ability of the compounds is tabulated in the above table. From
the table it is quite evident that compounds 1-4 having aliphatic groups in the amino acid part is
more susceptible to form gel with the paraffinic solvents. Another trend is observed that higher
the molecular weight of the paraffinic solvent, lower is the (Mininum Gelation Concentration)
MGC. Aromatic solvents are also effectively gelated following the trend that heavier molecules
gelated with more ease. Crude oil as well as different refinery distillates are also transformed to
gel by the gelator compounds. Minimum uptake capability of crude oil was found be from 40 to
50 times whereas, the maximum gelation ability is observed for diesel (76 to 100 times). From 1
to 4, the amino acid based alkyl chain length gradually decreases, which is responsible for the
decreasing order of their gelation ability (MGC for crude oil: 2 for 1 and 2.5 for 4). Gelation
ability of 1-4 with paraffinic solvents of the refinery distillates follows the trend as follows:
higher the amino acid chain length, lower is the MGC. Gelation ability of these compounds were
also tested for vegetable oil. All thegelators exhibited very effective gelation capability
comparable to that of diesel.
Table 1: Gelation abilities of 1-4 in different hydrocarbon solvents and oils
Compound
1
2
MUC
Hexane
MGC
(%w/v)
1.1
Octane
3
MUC
90
MGC
(%w/v)
1.1
90
MGC
(%w/v)
1.15
0.8
125
0.9
112
Dodecane
0.7
142
0.7
Hexadecane
0.6
166
Benzene
0.9
Toluene
Xylene
4
86
MGC
(%w/v)
1.1
0.9
111
0.9
111
135
0.8
133
0.7
135
0.7
138
0.7
142
0.8
123
117
0.9
117
0.9
112
0.9
111
0.7
142
0.8
125
0.8
119
0.9
111
0.7
142
0.9
117
0.8
121
0.9
111
CRN
1
100
1.2
83
1.2
83
1.4
71
SRN
0.8
125
1
100
1.1
90
1.1
90
Diesel
1
100
1.1
90
1.1
90
1.3
76
Crude oil
2
50
2.2
45
2.3
43
2.5
40
0.9
111
1.0
98
1.2
86
1.1
89
Vegetable oil
MUC
MUC
87
MGC = Minimum Gelation Concentration, MUC = Minimum Uptake Capability
In order to check the effect of the composition of crude oil on the gelation ability of the
organogelator, experiments were conducted with crudes with varying API gravities ranging from
low API (C1, 18.8⁰) to high API (C5, 40.5⁰). Highest gelation ability is exhibited by 1 and
decreasing order of gelation ability from 1 to 4 for other oils/solvents was exhibited. It could be
observed that heavy crude exhibited higher MGC and lighter crude had lower MGC and the
uptake capability decreased with increase in API gravity, thereby indicating the composition of
crude also played a major role in the uptake by the gelator.
Table 2: Gelation abilities of 1-4 in different crude oils
Compound
1
C1 (API = 18.8)
MGC
(%w/v)
2.2
C2 (API = 27.1)
2
MUC
3
MUC
4
46
MGC
(%w/v)
2.1
MUC
47
MGC
(%w/v)
2.1
MUC
47
MGC
(%w/v)
2.2
2
50
2.1
47
2.1
48
2.1
47
C3 (API = 28.1)
1.9
52
1.9
52
2.0
51
2.0
50
C4 (API = 35.5)
1.8
55
1.9
54
1.9
53
1.9
52
C5 (API = 40.5)
1.8
57
1.8
54
1.8
54
1.8
54
45
MGC = Minimum Gelation Concentration, MUC = Minimum Uptake Capability
The ease of formation of gel with lighter crude (C5) compared to the heavier crude (C1) could be
attributed to the higher paraffinic nature of the crude oil. Nevertheless, the gelators were efficient
for the most of the crudes covering the wide spectrum of crude basket available from different
parts of the globe. Phase selective gelation of petroleum products from a biphasic mixture of oil
and water was performed. Notably, all compounds were effective for oil phase gelation without
altering the water phase during performance evaluation. Compound wise gelation trend in the
biphasic mixture follow the same order as reported in Table 1 with minimal alteration to that of
individual oils. Maximum increment of MGCs for SRN, SRN and diesel were 0.2 % (w/v) from
their respective individual/single phase studies. Oil selective gelation of petroleum products from
a biphasic mixture of oil and sea water was also performed. Gelator compounds were able to
gelate exclusively the oil phase without gelating the sea water phase. Comparison of the results
led us to conclude that even under highly saline conditions, negligible changes/effect on gelation
properties (MCG & MUC) was observed. This observation indicated the strength and capability
of the organogelator towards the gelation preference for organic phase even under extreme
conditions. Their property could be tuned from hydrophobic to hydrophilic nature, i.e., the water
selective gelation ability, with a small change in the chemical structure. The gelator was
converted to its sodium salt which exhibited the reverse gelation property of capturing
water/aqueous phase in presence of hydrocarbon mixture.
Table 3: Gelation abilities of 1-4 in various oil-sea water mixture
Compound
1
2
MUC
Crude-Sea Water
MGC
(%w/v)
2
CRN-Sea Water
3
MUC
50
MGC
(%w/v)
2.3
1.2
83
SRN-Sea Water
1
Diesel-Sea water
1.2
4
MUC
43
MGC
(%w/v)
2.3
MUC
43
MGC
(%w/v)
2.6
1.3
76
1.3
76
1.4
71
100
1.2
83
1.2
83
1.3
76
83
1.2
83
1.3
76
1.4
71
MGC = Minimum Gelation Concentration, MUC = Minimum Uptake Capability
38
Capability of the gelators for practical application in oil spill recovery was tested by performing
a model oil spill set-up. As the prime focus was on oil spillage recovery, the process demands the
gelation at ambient temperature without any heating-cooling cycle on sea water. In the prototype
study, crude oil was added over water and gelator solution of 1 in hot toluene was added. Within
a few minutes, crude oil along with the aromatic solvent were transformed to the gel phase which
was collected by scooping out. The oil from the gel phase was recovered by vacuum distillation.
The recovery of oil from (crude oil) gel was 65% up to boiling point of 65% whereas, for diesel,
the recovery was > 95%.
Figure 2 Images of crude oil layer over water (A) before gelation and (B) after gelation.
Mechanical strength of the gels is an important parameter to collect gels over aqueous layer for
oil recovery. Stability of the gels was found to be very high, floating over the aqueous layer for
several weeks without deteriorating the gels. Both frequency sweep and amplitude dependent
rheology experiments exhibited typical viscoelastic gel behaviour by G' > G'' value. The storage
modulus value in the order of 104-105 Pa indicated a strong gel state to be retained when a small
strain was imposed. The strong mechanical gel strength was a consequence of three dimensional
fibrous networks as observed by FESEM studies. The three dimensional cross-linked fiber
lengths in gel was several micrometers having cross linking dimension of few nanometers as
formed in different solvents. Again, strong intermolecular hydrogen bonding, as revealed by FTIR study, was the driving force for three dimensional cross linking.
Figure 3 (A) Dynamic Rheology of the organogels obtained from different oils as a function of
angular frequency, (B) SEM images of xerogels obtained from diesel.
Conclusions
Four dipeptide-based compounds were synthesized with the variation of amino acid part in
between Methionine and Valine. Excellent gelation ability was exhibited by all four compounds
for hydrophobic solvents, several crude oils with varying APIs and refinery distillates (SRN,
CRN, Diesel) along with vegetable oil. Gelation ability of the compounds were found to follow
the trend of 1 > 2 > 3 > 4. Apart from oil phase gelation, structural modification to sodium salt
can exert water phase gelation too thereby extending the application of these gelators for the
removal of water from water-oil mixture where water is in lesser quantity
References
1. B. O. Okesola, D. K. Smith, Chem. Soc. Rev. 2016, 45, 4226.
2. On Scene Coordinator Report on Deepwater Horizon Oil Spill, 2011,
http://www.uscg.mil/foia/docs/dwh/fosc_dwh_report.pdf
3. Y. Gong, X. Zhao, Z. Cai, S. E. O’Reilly, X. Hao, D. Zhao, Mar. Pollut. Bull. 2014, 79, 16.
4. D. Dave, A. E. Ghaly, Am. J. Environ. Sci. 2011, 7, 423.
5. B. O. Okesola, D. K. Smith, Chem. Soc. Rev. 2016, 45, 4226.
6. K. Liu, P. He, Y. Fang, Sci. China: Chem. 2011, 54, 575.
7. Molecular Gels: Materials with Self-Assembled Fibrillar Networks, ed. R. G. Weiss and P.
Terech, Springer, Dordrecht, Netherlands, 2006.
8. P. Terech, R. G. Weiss, Chem. Rev. 1997, 97, 3133.
9. E. K. Johnson, D. J. Adams, P. J. Cameron, J. Mater. Chem. 2011, 21, 2024.
10. Ch. S. K. Raju, B. Pramanik, T. Kar, P. V. C. Rao, N. V. Choudary, R. Ravishankar, RSC
Adv. 2016, 6, 53415.
BackCasting-Using Integrated Planning
Back Casting – A novel Approach using
Integrated Planning
21st Refinery Technology Meet
20th to 22nd April 2017, visakhapatnam
Submitted byShri Ramandeep Singh,
Shri K. Mandal,
Shri S.N Pandey,
Deputy Manager (Optimization), IOCL, Corporate Office, New Delhi
General Manager (Optimization), IOCL, Corporate Office, New Delhi
Executive Director (Optimization), IOCL, Corporate Office, New Delhi
1
BackCasting-Using Integrated Planning
Table of Contents
Introduction ............................................................................................................................ 3
1.1 What is integrated planning .......................................................................................... 3
2 Back Casting ............................................................................................................................ 4
2.1 Operational variance Backcasting......................................................................................... 4
2.1.1 Variation in crude mix ................................................................................................... 4
2.1.2 GRM loss due to unplanned interruptions ..................................................................... 4
2.1.3 Variation due to Change in ISD ..................................................................................... 5
2.1.4 Loss due to variation in unit performance ..................................................................... 5
2.1.5 Actual Production vs. plan production........................................................................... 5
2.2 Integrated approach – Variance in boundary Inputs ............................................................. 6
2.2.1 Back Casting for Forecasted crude & product prices .................................................... 7
2.2.2 Back casting on Product Demand variance ................................................................... 8
2.2.3 Crude Mix variance........................................................................................................ 9
3. Conclusion ............................................................................................................................. 10
1
2
BackCasting-Using Integrated Planning
Back-casting: A novel approach using Integrated Planning
1 Introduction
IndianOil is India’s flagship Maharatna national oil company with business interests straddling
the entire hydrocarbon value chain – from refining, pipeline transportation and marketing of
petroleum products to Research & Development, Exploration & Production, marketing of natural
gas and petrochemicals. IndianOil Group (including two refineries of its subsidiary company
Chennai Petroleum Corporation Ltd. (CPCL)) owns and operates 11 of India’s 23 refineries with
largest and most extensive network of retail outlets with more than 24000 touch points.
Today, there are more opportunities for coordinating activities across a supply-chain even in
such complex operations as refining and Marketing, because of improving information systems
and communication technologies. For an organization like IndianOil that handles operations like
logistics, refining and marketing, an effective and integrated supply chain encircling crude
procurement, refineries production planning and products logistics management is required. The
intricacies of supply chain are being taken care through Integrated Planning (IP) involving linear
programming. It takes into account the diverse refinery configurations along with the constraints
involved in the entire supply chain. The allocation of crude to a particular refinery and the
positioning of a product in a desired location from the available sources and transportation mode
are taken care of by integrated planning model keeping in view the constraints involved, cost
incurred and the price realized. It thus optimizes the whole supply chain giving better margins.
1.1 What is integrated planning
The Integrated Planning Model (IP) is a combination of detailed refinery LP models of Barauni,
Gujarat, Haldia, Paradeep, Mathura, Panipat Refiney & petrochemical, Guwahati, Digboi &
BGR, refineries along with simple models (Extreme point) of CPCL & CBR refineries and all
India logistic networks for all the products. The IP model has been put in use since August 2004.
The Integrated Planning model has set a benchmark world over due to its size and complexity
which involves more than 80000 variables and 31000 constraints. The salient configuration
features of the model are as under:






Detailed LP models of Nine refineries and two extreme point refineries (IOCL subsidiaries
refineries- CPCL & CBR)
More than 160 crudes
800 demand locations after demand aggregation based on business logic
Crude and products pipelines along with pipeline Capacities and scheduling constraints
Four modes of transportation for products namely Pipeline, Rail, Tanker & Road
Provides an end to end integrated solution as a global margin considering all functions.
3
BackCasting-Using Integrated Planning
2 Back Casting
Backcasting is a performance management tool that provides insight as to where value is created
or where is it lost in the overall supply chain. Typically, Backcasting is a process to analyze the
variance in actual refinery operations from plan. Apart from identifying the overall gap in GRM
(gross refinery margins), purpose of Backcasting is to be able to classify the gaps into various
heads leading to reasoning & improvement actions to close the gaps. For example total variation
in GRM may be classified into following headers like change in Crude-Mix, loss due to
unplanned interruptions, loss due to change in ISD (intermediate stock depletion), variation in
unit performance & change in product mix. Such margin variance analysis is mostly focused on
calculating GRM loss due to operational variations.
For planning, Indian oil uses Integrated Planning Model (IP), a combination of detailed refinery
LP model of nine refineries and simple models (Extreme point) of CPCL & CBR refineries with
all India logistic networks for all the products. Such mammoth LP model involving more than
80000 variables and 31000 constraints creates challenges for traditional margin variation
analysis approach to categorise various gaps under different headers. To overcome this, Indian
oil has adopted two stage approaches to break up analysis into two parts namely operational
variance & boundary Input variance separately.
Operational variance mostly focuses on refinery standalone model to find out the impact of
operational factors like unit throughput, interruptions & intermediate stocks buildup/depletion
etc. Global factors like demand, prices & crude mix impact planning on corporate level and also
there is some extend of interaction between refineries for example reduction in demand of one
refinery may have shifted to another refinery. Therefore their impact is determined using
integrated planning to capture the global impact after adjustment of local variation.
Certain prominent losses like crude mix impact are calculated at both standalone & corporate
level for deeper understanding of its variation. These variation can individually summed up again
at the corporate level to confirm the results of integrated planning backcasting.
2.1 Operational variance Backcasting
Operational variance is categorized and calculated in typical refinery sub models to categories
various refinery operational deviation like
2.1.1 Variation in crude mix
In the plan, certain crudes are planned to be available for processing during the month.
However due to tanker slippage or force majeure, the actual crudes processed during the
period may be different from those planned, having certain impact on GRM. This step of
backcasting tries to quantify the losses that arise due to change in crude mix arising out
due to crude scheduling issues like delay in crude cargo arrival, crude pumping schedule
from the ports etc.
2.1.2 GRM loss due to unplanned interruptions
This step of Backcasting aims at capturing the loss in GRM due to the unit unavailability
due to unforeseen interruptions.
4
BackCasting-Using Integrated Planning
2.1.3 Variation due to Change in ISD
This step of Backcasting aims at capturing the loss in GRM due to the change in ISD if
any applicable due to factors other than units’ interruption.
2.1.4 Loss due to variation in unit performance
This step aims at capturing the GRM loss due to changes in unit yield and catalyst
performance (eg operational inefficiency due to equipment constrains etc) as compared to
the plan.
2.1.5 Actual production vs. plan production
This will help in capturing GRM loss due to deviation in production plan due to ullage
constraints and actual product off-take resulting in changes in product mix and ISD
depletion / built-up not considered in the plan
5
BackCasting-Using Integrated Planning
2.2 Integrated approach – Variance in boundary Inputs
Integrated Planning approach utilizes the synergies that exist among crude procurement, refining
and product marketing to maximize the corporate profit. Given the large volumes that are
involved, small percentage changes translate into huge benefits to the Corporation.
Integrated planning model planning boundary starts from crude supply points for feeds and uses
depots prices & demand for various products. There are no refinery standalone inputs like
refinery transfer price or local product demands as illustrated in fig below.
6
BackCasting-Using Integrated Planning
The Integrated Planning model has following inputs:









Crude availability at ports
Grade wise products Demand at locations
Committed exports / Imports
Exchanges with Other Marketing Companies (OMCs)
Crude landed cost at ports
Products prices at demand location
Desired Inventory buildup / depletion at depots level
Exchanges with Other Marketing Companies (OMCs)
Planned refineries process units shutdown schedule
Theses inputs can broadly be categorized under three major areas which could be back-casted for
improved planning process.
a) Forecasted Prices
b) Forecasted Demand
c) Forecasted Crude Mix
2.2.1 Back Casting for Forecasted crude & product prices
For crude procurement & production planning integrated planning uses Crude & product prices that
are build on forecast quotes. Forecasting for prices is done four months in advance and may tend to
change during actual processing month. This volatility of prices may lead to different loss in
margins on account of change in product cracks or sub-optimal planning. The impact of price
volatility on margins is accessed by back-casting of integrated planning model on actual price basis.
7
BackCasting-Using Integrated Planning
Backcasting of forecasted prices help us to understand the gap between plan & actual quotes so
that we can improve upon the forecasting methodology. Also, in order to protect our refinery
margins, we have an option to hedge our margins in forward’s market. We can hedge most of the
major variables like Brent & Dubai price, products prices, brent & Dubai spreads etc and thereby
providing opportunity to protect & improve margins.
2.2.2 Back casting on Product Demand variance
IndianOil operates the most vast & extensive Hydrocarbon network of refineries, Terminals and
Depots. Multiple locations with multiple mode of transportation create an opportunity to globally
optimize demand allocation to refineries. This unique product slate creates a huge challenge for
forecasting of product demand numbers with their respective linkages to various source points.
8
BackCasting-Using Integrated Planning
Multiple mode of
transportation
Multiple source
& destinations
Various planning activity for IndianOil like production planning, logistics planning, contingency
planning etc require demand forecasting. All these activities are carried out for primary distribution,
i.e. for positioning of product upto depot/terminal level. One of the vital inputs to this entire process
is "Location wise product wise demand forecast". The process adopted is explained with figure
below as per hierarchy.
Deviation of forecasted demand to actual demand results in product inventory variation which in
turn impact refinery crude processing capacities. For example, import/export are planned one
month in advance with forecasted demand, any change in forecast demand will either results in
imports of products in deficit scenario or Export / Refinery throughput cut in case of surplus
scenario. Typical variation in planned demand to actual demand for major product like MS &
HSD is plotted below.
9
BackCasting-Using Integrated Planning
Back-casting of demand variation not only helps up to improve forecast but also to understand
inventory management capabilities & limitations in order to take care of small unexpected
variations.
2.2.3 Crude Mix variance
Crude oil forms the vital part of any refinery operation. Hence, selection of the proper crudes is
the foundation for smooth refinery operations, improving margins and delivering quality
products. With the wide range of crudes each having different properties and yields, it becomes
very important to make the right selection of crudes to feed the appetite of our refineries. Crude
mix planned Vs actual impacts refinery crude processing throughput, Secondary unit capacities,
product slate and hence has huge impact on GRM.
All major IOCL refineries share crude pipeline with other refineries. These crude pipelines limit the
Quantity & Quality of crude mix processing to specific requirements of refineries. If individual
refineries plan their crudes on standalone basis then it may result in unique crude grade & quantity
requirement, which may not be practically feasible to pump through pipelines. Therefore it is
pertinent to analyze the impact of crude mix on integrated planning model which will give holistic
loss / gain in margins.
3. Conclusion
The paper presents the scope of profitability improvement by carrying out back-casting studies
on integrated planning model. Back-casting on IP model will provide us with various areas
which can be independently accounted & improved for margin gain.
The figure below represents the journey of planned GRM to actual GRM due to various factors
which can be individually classified & improved.
10
BackCasting-Using Integrated Planning
11
Panipat Refinery
21st Refinery Technology Meet, Vishakapatnam (20-22 Apr,2017)
Replacement of PTA Slurry Incenerator by Envirnomental Friendly Flaker at Panipat Refinery
 Introduction:- Slip stream of recycling Mother liquor(Containing Acetic acid, Corrossion products,
Manganese Acetate and cobalt Acetate) ,after removal of Acetic Acid is sent to incenerator, wherein slurry
is burnt off. The ash so generated is e-auctioned to parties.
 Comparison of Incenerator with Flaker:Sr no
Incenerator
1
Energy intensive
Flaker
Less Energy intensive
2
Emits Sox and Nox to Atm.
Environmental friendly
option.
3
Only Metals can be recovered
after Inceneration
4
5
High operation cost
Product less valuable
Metals as well as Organic
acids can be recovered from
Flakes.
Less Operation cost
Products are more valuable
 Salient Feature of a Flaker
 Feed in the form of melt / molten mass is applied as a thin film over the rotating metallic cylinder.
 The feed is introduced in a feed trough, provided with a agitator to avoid setting of the feed
material and for enabling uniform contact of feed slurry with the outer surface of the flaker drum .
 Cooling of the drum surface is achieved with the help of water spray jets from inside. The water jets
hit the inner shell at high velocity which results in effective heat transfer.
 The drum is kept rotating with the direction of rotation being towards the blade.
 The drum picks up material film from the feed trough & gets cooled during its passage to the
scraper blade.
 The solidified film is scraped off by the scraper blade to yield cooled flakes .
 The flaker drum enclosure is provided with suitable steam tracing arrangement to avoid vapour
condensation and to avoid deposition of benzoic acid inside the flaker enclosure. Enclosure will
be fume and dust tight with provision of inert gas purging
 Suitable vapour carrier system with wet scrubber and ID fan is provided for trapping benzoic acid
fumes if any.
Indicatve diagram of Flaker
Two storied View of Flaker
Ist Floor View of Flaker
Bottoms chutes for Flaked Product
Paper for RTM 20–22 April 2017, Visakhapatnam, India
CORROSION MITIGATION IN HYDROTREATERS BY EMPLOYING
BEST DESIGN & OPERATING PRACTICES
AUTHORS:
Rakesh Gagat (Deputy Process Manager, RHQ (T), Indian Oil Corporation Limited, New Delhi)
Mukesh Kumar Sharma (Chief Manager, RHQ (T), Indian Oil Corporation Limited, New Delhi)
1. INTRODUCTION
Refineries today are facing new challenges in order to meet the stringent fuel
specification and emission norms that imply huge investments in refineries.
India is in critical phase of rapid transition from BS-III to BS-IV automotive
fuels specifications (by April 2017) and subsequently to BS-VI norms in short
span of next three years. Both the primary transportation fuel, Ms & diesel
should conform to sulfur level of 50ppm & 10ppm in BS-IV & BS-VI standards
respectively.
In meeting such stringent sulfur specification, hydrotreating technologies are
gaining more & more importance. This sets new challenges to further develop
processes/designs of hydrotreaters which are more reliable, energy efficient,
safe to operate and in turn meet the plant-run time targets. Refineries are
increasing plant run time targets, since downtime means process losses and
incur cost including maintenance and safety costs.
Diesel Hydrotreater unit (DHDT) in one of the Indian Oil refinery was facing
severe corrosion downstream of high pressure (HP) separator drum, stripper
feed circuit and stripper column top trays section since commissioning of the
unit. Several mitigating measures were employed such as intermittent water
wash, feed bypass arrangement across reactor feed/effluent exchangers,
stripper internal metallurgy up-gradation etc. without much success.
Category: Operational Excellence and Best Practices
1
Paper for RTM 20–22 April 2017, Visakhapatnam, India
Process Design Engineering cell of Indian Oil Corporation Limited has carried
out rigorous study to analyze root cause of perennial corrosion problem and
suggested a design change to completely mitigate the issue.
2. OBJECTIVE
 Corrosion mitigation encountered due to ammonium salts in Indian Oil
Corporation Limited (IOCL) DHDT Unit
3. UNIT CONFIGURATION (EXISTING)
Figure 3.1: Unit Configuration (EXISTING)
Category: Operational Excellence and Best Practices
2
Paper for RTM 20–22 April 2017, Visakhapatnam, India
Unit Description
Feed / Reaction / HP Separation section
The blended Diesel stream is filtered through the feed filter and sent to the feed
surge drum. The feed is pumped under flow control to the two parallel reaction
section trains, via the feed pump. Antifouling agent diluted in hydrotreated
diesel is injected in the feed upstream of the feed pump.
The hydrogen make-up gas coming from battery limits is routed to the makeup compressor KO drum. From there, it is compressed by the make-up
compressors. The compressed gas is sent to the discharge of the recycle
compressor to join the recycle gas and to be used as cold quenches and recycle
gas. Each reaction section train is made of reactor feed/effluent exchanger,
stripper feed preheater, reactor heater and two reactors, and can process 50%
of the total feed.
The pumped feed is split into each reaction section train and then mixed with
make-up/recycle hydrogen stream. The mix is also preheated in the reactor
feed/effluent exchangers. The reactor feed is then brought up to the required
temperature in the reactor heaters. The reactor inlet temperature is controlled
by acting on the fuel oil/fuel gas burners.
The fluid is routed to the first reactor (HDS section), and includes three beds in
order to limit the temperature increase inside the reactor. Cold quenches are
injected between the beds under flow control reset by catalyst bed inlet
temperature. The HDS reactor effluent is quenched and sent to the second
reactor (HDT Section), and includes two beds. A cold quench is injected
between the beds under flow control reset by catalyst bed inlet temperature.
The HDT reactor effluent is used to exchange heat with the stripper feed in the
stripper feed pre-heater under temperature control of the stripper feed. The
Category: Operational Excellence and Best Practices
3
Paper for RTM 20–22 April 2017, Visakhapatnam, India
effluent outlet stream of this exchanger is then used to preheat the reactor
feed. The cooled effluents of each train are mixed together and then collected in
the hot HP separator. As indicated previously, the temperature of this drum is
controlled by adjusting the amount of feed by-passed.
The vapor phase from the hot HP separator is routed to cold HP separator.
MP Separation / Stripper section
In the cold MP separator, three phases are separated. The vapor phase is sent
to MP separators vapor K.O. drum under pressure control together with the hot
MP separator vapor. The hydrocarbon liquid from the cold MP separator is
withdrawn under level control, mixed with the hot MP separator liquid and
routed to the stripper section.
At the hot MP separator, the vapor phase is mixed to the cold MP separator
vapor under pressure control. The mix is cooled down in MP separators vapor
water cooler and sent to MP separators vapor K.O. drum. The off gas is then
sent directly at battery limit to OHCU and the condensate is sent under level
control to stripper reflux drum.
The hydrocarbon liquid phase from the hot MP separator is mixed under
differential pressure control with the hydrocarbon liquid from the cold MP
separator. The mix is routed to the stripper feed/bottom exchanger for heat
recovery and then split into two identical streams. Each stream is under flow
control reset by the hot MP separator level control. The required stripper inlet
temperature is obtained by heat exchange in-between these two stripper feed
streams and the reactor effluent of each train in the stripper feed pre-heaters.
Medium pressure superheated steam is injected under flow control at the
bottom of the stripper in order to obtain a hydrotreated diesel with the correct
flash point and free of H2S.
Category: Operational Excellence and Best Practices
4
Paper for RTM 20–22 April 2017, Visakhapatnam, India
Stripper overhead is partially condensed first in the stripper air condenser and
then in the stripper trim condenser. Corrosion inhibitor diluted in stabilized
naphtha is injected into the column overheads to protect the condensers.
The stripper bottom product is pumped by the hydrotreated diesel pump and
exchange heat against the stripper fee. The dry hydrotreated diesel product is
finally sent under stripper bottom level control to battery limit.
4. CORROSION STUDY
It was informed that DHDT unit is facing severe corrosions in stripper feed
circuit and Stripper Column Top Trays section. The same was referred to
Process Design Engineering Cell (PDEC), IOCL to study. PDEC carried out the
detailed study after rating of the existing units.
OBSERVATIONS OF THE STUDY
The likelihood of deposition of NH4HS & NH4Cl at the design level of H2S and
NH3 at Packinox outlet temperature of 180 Deg C was checked using API 932-B
Deposition Charts.
The observations are as under:
1. NH4HS can never deposit at that temperature/ pressure (The deposition
temperature is well below 100 Deg C).
2. There is high likelihood of deposition of NH4Cl if HCl content in packinox
vapor goes even beyond 1ppm at 180 Deg C.
Category: Operational Excellence and Best Practices
5
Paper for RTM 20–22 April 2017, Visakhapatnam, India
FURTHER OBSERVATIONS OF THE STUDY:

It has been also observed that Providing water upstream of Packinox
exchanger may help in preventing NH4Cl deposition but cannot eliminate
the

problem.
At Packinox outlet condition it has been observed that most of the water
remains in vapor phase (around 90-95 %) and hence sufficient water may
not be available in aqueous phase to absorb NH4Cl / HCl.

Thus, solid NH4Cl may remain with hydrocarbon streams from the HHPS
bottom & gets fed to stripper after heating.

The feed NH4Cl may resublime in NH3 and HCl at stripper inlet
temperature of 255 deg C and may redeposit at the lower stripper top
temperature of 150 Deg C. This may give HCl corrosion in the feed zone and
above.
Solution/Modifications Proposed :

HHPS to be operated at 235-240 Deg C to reduce the deposition of
ammonium chloride salt at the outlet. The Rx effluent temperature at the
outlet of Packinox can be maintained through bypass of feed if required.

One New exchanger in HHPS vapor (HM-01) (HP/HP) is proposed to heat the
feed prior to its entry into Packinox-1.

One New exchanger in HLPS Liquid (HM-02) (LP/HP) is proposed to heat the
feed prior to its entry into Packinox-2.

Recycle gas will be mixed with the feed as existing at the Packinox inlet.
Category: Operational Excellence and Best Practices
6
Paper for RTM 20–22 April 2017, Visakhapatnam, India

Increasing the temperature of HHPS to 235-240 Deg C from present level of
around 180 Deg C shall result in higher vapor flow through air cooler
resulting in increased flow and duty.

Keeping the same air cooler duty as given in process package (Case-2 EOR),
this shall manifest in higher CHPS temperature (70 0C) and increased vapor
flow to HP amine absorber. The condensed hydrocarbons will be returned to
CLPS vessel thereby reducing load to amine absorber and RGC.

In order to reduce the temperature and part load at the inlet of HP amine
absorber, a new water cooler is proposed.

The new water cooler may be necessary to contain the capacity of RGC and
also to take care of operation beyond design capacity.
5. UNIT CONFIGURATION (MODIFIED)
The Modified unit Flow scheme is shown in the figure 5.1 below.
Category: Operational Excellence and Best Practices
7
Paper for RTM 20–22 April 2017, Visakhapatnam, India
Figure 5.1: Unit Configuration (MODIFIED)
6. CONCLUSION
PDEC, IOCL recommended that HHPS is to be operated at 235-240 Deg C to
reduce the deposition of ammonium chloride salt at the outlet. Later the
process licensor of the unit was awarded the job for detailed study for
mitigation of corrosion issue and suggested modifications were almost in line
with the PDEC, IOCL study.
Thus a good process design & operating practices should consider a holistic
approach covering all aspects like energy efficiency, reliability, runtime and
feed quality with all probable impurities.
Category: Operational Excellence and Best Practices
8
MANAGING REDOIL FOULING IN DEPROPANISER COLUMN
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P. KAMATH
PROCESS ENGINEERING DEPT., MRPL
ABSTRACT
column run time with varied operating conditions,
Petro Fluid Catalytic Cracking unit (PFCC)
mitigation of column fouling and recovery of the
employs deep catalytic cracking technology to have
fouled column, along with options for prevention of
selectivity towards olefins maximization. Increased
Depropanizer column fouling.
selectivity of the catalyst towards olefins will result
Keyword: PFCC, Debutanizer, LPG, carbonyl
in the generation of aldehyde/carbonyl components
compound, Caustic Treatment, Redoil Polymers.
as a byproduct. These components have a tendency
to polymerize in alkaline medium and form
Introduction
agglomerates typically known as Redoil which is
Petro Fluid catalytic cracking (PFCC) utilizes
nothing but a red color polymer formed because of
propylene maximization technology, along with the
Aldol condensation reactions.
generation of Gasoline and LPG.
Depropanizer is
Depropanizer is used in FCC units to
used in PFCC units to separate Treated/Sweet LPG
separate Treated/Sweet LPG feed to C4 LPG and
feed to C4 LPG and C3 LPG. C3 LPG is further
C3 LPG. C3 LPG is a valuable product for refiners
treaded and separated for Propylene recovery in
due to good Propylene content. Because of the
Propylene
presence of mercaptan (RSH) in sour LPG, Caustic
Depropanizer column is sour LPG.
treatment is preferred upstream of Depropanizer
contains impurities like H2S, mercaptan (R-SH),
column. However presence of carbonyl compound
CO2 & carbonyl etc. Removal of these impurities is
in the Sour LPG stream creates an ideal condition
important for success of LPG copper corrosion test
for Aldol condensation reaction with Na+ ions as
and on spec feed for Propylene Recovery Unit
catalytic promoter, resulting in the formation of
(PRU). As typical Amine Treatment is not effective
polymer. Red Oil creates fouling issues in
for removal of RSH and mitigating hydrogen
downstream section of caustic treatment. There is a
sulphide to meet copper strip corrosion test
good probability for possible caustic carry over
requirement, caustic treatment of sour LPG is
along with redoil to Depropanizer column, with any
preferred.
disturbance in the normal plant operation or
Carbonyl compounds (like Acetaldehyde) forms as
fouled/damaged filter/coalescer after LPG water
a byproduct of catalytic cracking reaction. Presence
wash System. Once caustic along with redoil
of carbonyl in sour LPG can be checked using
polymer enters the Depropanizer column, it tends
ASTM D4423-00. A measured amount of sample is
to settle in feed flash zone area and reduces the
reacted
efficiency of separation of C4 LPG and C3 LPG
hydrochloride
drastically due to fouling of tray & downcommer
amount of hydrochloric acid. Results are reported
leading proportional reduction in mass transfer area
as milligram per kilogram carbonyl as acetaldehyde
for effective separation. This paper intends to bring
[1]. PFCC sour LPG reported 75 mg/Kg carbonyl
out the practical experience of behavior of fouled
as acetaldehyde which is substantial amount to
Depropanizer column because of redoil formation
initiate carbonyl Aldol polymerization during pre-
in the feed flash zone, techniques of prolonging the
treatment of sour LPG. Carbonyl compounds get
Recovery
with
Unit
alcoholic
solutions,
(PRU).
Feed
to
Sour LPG
hydroxylamine
releasing
equivalent
1
MANAGING REDOIL FOULING IN DEPROPANISER COLUMN
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P. KAMATH
PROCESS ENGINEERING DEPT., MRPL
absorbed into the caustic and through aldol
The rate of Aldol condensation reaction increases
condensation
with temperature acetaldehyde concentration and
reaction
mechanism,
forms
hydrocarbon Polymers. Below mentioned equation
shows
Aldol
condensation
caustic concentration.
mechanism,
As a common practice caustic strength is usually
being maintained on a bit higher side to overcome
any the possible sour gas slippage to downstream
Polymer unit. However this increases the potential
for Aldol condensation reactions
Red oil Polymer formation is the most common
cause
of
fouling
in
equipments
like
LPG
filter/coalescer spent caustic Pump strainers and
even in Wet Air Oxidizer (WAO) F/E exchangers
choking and plugging[3] etc. High DP across LPG
coalescer is clear symptom of red oil fouling at
LPG coalescer and this could occur after pre filter
damage. If coalescer is not rectified caustic
Fig1: Red Oil formation mechanism
carryover to Depropanizer is most probable due to
The oily polymer appears red so the usually
inefficient coalescing of caustic. Above situation
referred to as “red oil”. Red colour is caused by the
creates recovery issues of valuable streams or
number of double bonds in the reaction products
unscheduled plant turnaround.
[2]. The initial, smaller polymer remains soluble in
the caustic. As the polymer grows, it forms light
insoluble oil that floats on top of the aqueous
phase. It will continue to react and form heavier
polymers, eventually forming solids in the caustic.
The red oil will absorb other organics from the
cracked gas as well as corrosion products, and thus
increase in volume.
Background
Practical experience of behaviour of Fouled
Depropanizer Column
In PFCC unit High Delta Pressure across Sweet
LPG Prefilters were observed therefore standby
Prefilters were taken in line and exhausted Prefilter
candles were replaced, PreFilters candles found
damaged and with red oil sludge deposited and this
may have lead to the Slippage of red oil particle to
Coalescer as after some time high Coalescer DP
was observed. High Delta pressure across the
Coalescer affected the performance of Coalescer
and reduces the water/caustic coalescing result in
the caustic carryover to downstream depropanizer
Fig 2: Red Oil sludge
column. Coalescing efficiency can be analyzed by
2
MANAGING REDOIL FOULING IN DEPROPANISER COLUMN
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P. KAMATH
PROCESS ENGINEERING DEPT., MRPL
measuring pH and presence of Na+ ions in. high
pH as 11-12 indicates inefficient coalescing.
2.00
LPG FILTERS
1.50
350
0.6
300
0.5
250
LPG COALESCER
0.4
200
1.00
0.3
150
0.50
100
50
0.00
0.2
Avg Reflux
Avg DP
0
Fig 3: High DP across the coalescer
Continuous Flooding conditions were observed in
0.1
0
Fig4: Depropanizer column operation at lower
reflux
the Depropanizer column rectification section as
Based on primary analysis to overcome fouling of
Delta pressure reached up to 0.8 Kg/cm2. Decrease
Depropanizer column, water washing of the
in efficiency of the Depropanizer Column tray was
column was proposed without taking entire unit
observed as In flooding condition C4 in PRU feed
shutdown. It was planned to route entire Treated
which is the top product in the Depropanizer
Propylene rich LPG to rundown without recovering
column used to go as high as 4~5wt% against the
Propylene and isolating Depropanizer column, also
design
Downstream PRU section was kept in circulation
of
0.5wt%,
propylene
slip
from
Depropanizer bottom is as high as 9~10wt%, which
mode for quick in lining of PRU section.
is equivalent of 5.5~6 TPH of Propylene loss in
Washing arrangements includes supply hot water
LPG. Depropanizer top temperature used to go as
from Depropanizer top (Reflux drum) and draining
high as 50OC against the design of 46OC due to
arrangements from bottom section.
heavier component accumulation. Increase in reflux
flow had very little impact on top temperature
control.
Techniques of prolonging the column run time
with varied operating conditions
The flooding condition was normalized by reducing
the reflux and reboiling. Column was operating
under reduced load conditions by forgoing the
product quality to allowable limits. However
frequent column flooding was observed.
Fig5: Depropanizer washing arrangement
3
MANAGING REDOIL FOULING IN DEPROPANISER COLUMN
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P. KAMATH
PROCESS ENGINEERING DEPT., MRPL
Initial pH test of drained liquid found in strong
Parameters
Sample 1
Sample 2
basic range (pH 12) and colour was blackish, this
LOI @ 900C
80.5%
95.4%
conformed both the presence of caustic in column
Toluene
Insoluble
Insoluble
and washing arrangement was efficiently working.
solubility/color
/yellow -orange
/yellow -orange
Iron as
3% & 4.3%
0.28%/1.12%
Fe/Fe2O3
With continuous washing for 24 hr, slowly
improvement in both pH and colour was observed.
After normalizing the wash water pH to 7 and
observing that wash water colour appeared clean,
Depropanizer water wash was successfully called
FIG 6: Drained liquid form Debutanizer
off. After the re start-up Depropanizer column
Depropanizer reboilers bottom were also drained
to OWS. LPG filter and Coalescer were replaced
performance was observed to be normal.
Conclusion
with new candle.
Red oil formation is inherently associated with
caustic and olefin system.
Proper monitoring and primitive action can avoid
unplanned shutdown due to red oil fouling
FIG 5: LPG filter and Coalescer with red oil
fouling.
Red-yellow
colour
mucks
was
found
in
Depropanizer bottom pump strainers. Lab analysis
of these samples shows a high LOI with presence
of iron rust as red oil will absorb other organics
from the cracked gas as well as corrosion products,
that muck was polymeric hydrocarbon same is
proved with insolubility in toluene and producing a
separate
orange-yellow
colour
layer

Maintaining Caustic strength in
permissible range.

Frequent sampling of Coalescer out let
should be carried out to check Coalescer
efficiency.

Filters and Coalescer should be
immediately rectified if high Delta
pressure is observed

In case frequent fouling issues trial of
commercial red oil inhibiters and changing
the normal Coalescer to more efficient LL Coalescer can be considered
Water wash is a good option to recover the fouled
column without opening and taking entire unit
shutdown.
References
when
dissolved in toluene.
4
MANAGING REDOIL FOULING IN DEPROPANISER COLUMN
MANNU JHA, NAGARAMESHKUMAR PARIMI, ANAND A. HARADI, SHYAM P. KAMATH
PROCESS ENGINEERING DEPT., MRPL
1.
2.
3.
ASTM D4423-00, Standard Test Method
for Determination of Carbonyls in C4
Hydrocabons.
Blaschke, M, ‘Caustic tower fouling
:identifying the causes” American Institute
of chemical engineers, 15th Ethylene
producers conference ,New Orleans,
lousisiana, March 30-April 3,2003.
Clayton M, the effect of caustic tower
operation and spent caustic handeling on
the zimpro WAO of ethylene spent
caustic,Ethylene production conference
April 2009, Tampla,FL
5
1
21st Refinery Technology Meet (RTM)
STEAM COIL BYPASS IN OUT BOARD STEAM GENERATOR IN CCRU
G L Arun Kumar, Appalaraju Pentakota, V.Venkatesh, S.Sankaran
ABSTRACT:
The reformate stream requirements of Motor Spirit (Petrol) pool in Chennai Petroleum
Corporation Limited (CPCL) is met with a 0.3 MMTPA CCRU unit, which produces reformate of
102 RON at an average rate of 43 m3/hr. Apart from reformate the unit also exports HP steam
at the rate of 13.8 TPH and supplies 17000 nm3/hr off gas to PSA II which in turn contributes
99.99% pure hydrogen to the header at the rate of 12000 nm3/hr.
As per the original design, Flue gas from CCRU inter heaters (8000C) are collectively
taken to an Out Board Steam Generator (OBSG) which is a waste heat boiler to produce HP
steam (32 kg/cm2, 3100C). Flue gas temperature effectively drops from 8000C, at various
stages in OBSG, to 1800C, before being let out to the atmosphere. Boiler Feed Water (BFW, 43
kg/cm2, 970C), on the other hand, passes through a pre heat coil inside the steam drum and
then enters the Economiser section of OBSG at 2200C. Flue gas (2600C, after heat recovery in
various other stages of OBSG) is allowed to exchange heat with this pre heated BFW in
Economiser to increase its temperature to 2400C. This pre heated BFW is then used for the
production of HP steam in OBSG a rate of 11TPH.
In 2014, the throughput of CCRU was increased from a design capacity of 52m3/hr to 60 m3/hr,
by adopting various in house de-bottlenecking procedures. As a consequence of this capacity
increase, the load on the inter heaters increased and this in turn increased the OBSG outlet flue
gas temperature to 2200C.
This paper deals with this particular problem of high flue gas outlet temperature and how
this was seen as an opportunity to increase steam production in OBSG. Increased heat
recovery in economiser was seen as a solution to decrease the flue gas outlet temperature. In
order to increase the heat recovery, the difference in temperatures of the BFW and flue gas had
to be widened. As a solution, part of BFW to steam drum coil was allowed to bypass the heat
recovery coil inside the steam drum to enter the economiser at lower temperature. This
increases the temperature difference between BFW and flue gas and increases the quantum of
heat recovered from the flue gas, thereby decreasing the outlet flue gas temperature.
2
Bypassing about 3 TPH of BFW increased the approach by about 800C and the flue gas
temperature dropped to 180oC. This again translated into an increase in HP Steam production
of about 2 TPH.
1. PROCESS DESCRIPTION:
The OBSG system has been designed to maximize energy recovery from the outgoing flue
gases of the Reformer heaters (F3/F4/F5/F110) & stabilizer reboiler heater (F-106). It is
designed to produce 13.8 TPH of steam from the heat available in the flue gases.
Waste Heat Recovery system (OBSG) consists of a single drum natural circulation waste
heat boiler followed by an ID fan and stack. ID fan draws the flue gases from the common duct
connecting the stack of all the furnaces through OBSG. There is one bypass duct at the
common stack to divide between the incoming flue gas to ID suction & the flue gas from the
discharge of ID fan. There are studded tubes in OBSG, which provides extended surface area
for better heat transfer between DM water or stream & flue gases. The flue gases from the
process pass through the hot gas ducting and enter the waste heat boiler. The gases are cooled
from about 821°C to about 160 -225°C. Then they are passed through an ID fan to the exhaust
stack.
BFW on the other hand enter the system from steam drum coil (98°C ) at the
economiser and then passes through pre evaporator, evaporator and finally to the super heater
section and is then put to the HP steam header after de superheating. The temperature profile
of OBSG is shown in Fig 1.
OBSG consists of the following parts:
I.
Pre-evaporator:
The pre-evaporator tubes are of bare construction. These tubes act as screen tubes to protect
the super heater tubes from heat fluctuation. This section has 4 tubes along the gas flow with an
effective heat transfer area of 38.54 m2.
II.
Super heater bank:
The super heater is a counter current type heat exchanger with gas flowing vertically
downwards and the heat transfer tubes arranged horizontally. A de-super heater is placed at the
outlet of super heater to achieve the required temperature of super heated steam. This section
has 4 tubes along the gas flow with an effective heat transfer area of 38.54 m 2.
3
III.
Evaporator Bank:
The evaporator bundle of forced circulation type is situated below the super heater. The gas
flow shall be vertically downwards as in the case of super heater and the heat transfer tubes
shall be arranged horizontally. The entire evaporator bank is in turn is connected by riser down
comer circuit to the main steam drum. Suitable pitching of tubes is provided to suit the gas side
temperature profile. This section has 26 tubes along the gas flow with an effective heat transfer
area of 630.24 m2.
IV.
Economizer Bank:
The economizers are placed below the evaporator. This shall consist of horizontal heat transfer
tubes, with the gas flowing vertically downwards. This section has 16 tubes along the gas flow
with an effective heat transfer area of 287.64 m2.
V.
Boiler steam drum:
The boiler unit has been provided with one steam drum of welded construction. The
Drum contains a set of internal piping for feed water, dosing chemical distribution and a
demister, which will be capable of providing steam of the purity. The steam drum also
has an in built heat recovery coil through which BFW passes before entering the
economiser section.
2. BASIS OF DESIGN:
The inter heaters in CCRU unit are used to heat the process stream to the required temperature
to drive the favorable reaction in the reactors and hence achieve the required product quality.
These inter heaters in turn supply flue gas to the waste heat boiler to produce steam. The
typical characteristics of the flue gas stream and steam produced are given below.
a. Gas Data:
Gas Flow Rate
170 T/hr
Gas Temperature at inlet
821 °C
Gas Temperature at outlet 225/180 °C
Gas pressure
-25 mm WC (at the inlet of
the boiler)
Dust loading
250 mg/Nm
Gas side pressure drop
100 mm WC
Table 1. Gas data
4
c. Steam Data:
Pressure (min./nor. /max)
29.5/30.5/32.5 kg/cm2
Temperature
(min./nor.
27O/28O/290 °C
/max)
Steam Generation
13.800 Kg/hr
Water flow rate
14.215 Kg/hr
Feed water temp, at
105/110 °C
Battery Limit
Table 2. Steam data
3. TEMPERATURE PROFILE:
The temperature profile in the Waste Heat Recovery System is as shown below. The
bottom line shows the flue gas path. The flue gas is cooled from around 82I°C to around 225°C
in the system. The bottom line shows the steam/water line. The Boiler feed water enters the unit
at 105°C. The superheated steam leaves the unit at around 295°C.
The temperature profile of OBSG is shown in Fig 1.
800
755
700
295
240
240
260
220
220
105
SUPER HEATER
EVAPORATOR
PRE EVAPORATOR
ECONOMISER
DRUM COIL
Fig 1.Temperature gradient in OBSG: Before steam coil bypass
5
4. SCOPE FOR IMPROVEMENT:
The throughput of CCRU was increased from a design capacity of 52m3/hr to 60 m3/hr,
by adopting various in house de-bottlenecking procedures. As a consequence of this capacity
increase, the load on the inter heaters increased and this in turn increased the OBSG outlet flue
gas temperature to up to 2200C. This provided a scope for improvement in this section. High
flue gas temperature translated to possibility of increasing heat recovery from the flue gas.
Fig 2. OBSG Original Condition
6
5. SCHEME DESCRIPTION AND IMPLEMENTATION:
As a consequence, a study was taken up to increase the heat recovery from the flue
gas. The temperature profile, as shown in fig 1 , showed that the difference in temperature of
flue gas inlet and BFW inlet in the economiser section is only about 40 0C. Low delta T between
the hot and the cold fluids mean low heat recovery.
In order to increase the heat recovery, the difference in temperatures of the BFW and
flue gas had to be widened. So, part of BFW to steam drum coil was allowed to bypass the heat
recovery coil inside the steam drum, as shown in fig 3, to enter the economiser at lower
temperature. This increases the temperature difference between BFW and flue gas and
increases the quantum of heat recovered from the flue gas, thereby decreasing the outlet flue
gas temperature to 1800C.
3 TPH
13 TPH
180 C
260 C
80 C
DELTA T
180 C
Fig 3. OBSG Current status
7
Flue gas
800
BFW /Steam
755
700
295
240
240
260
180
SUPER HEATER
EVAPORATOR
180
105
PRE EVAPORATOR ECONOMISER
DRUM COIL
Fig 4.Temprature gradient in OBSG: After steam coil bypass
6. RESULTS AND DISCUSSION:
The effect of this bypass flow on various parameters is shown in table 8.
Parameter
Units
Mass Flow of Flue gas
T/hr
cP of flue gas
kcal/kg/K
Flue gas inlet temperature
Original condition
Present condition
150.00
169.50
0.24
0.24
C
780.00
800.00
Flue gas outlet temperature
C
220
185.00
wHeat removed from Flue gas
Mkcal/hr
20.23
25.10
Heat absorbed @ 67 % efficiency
Mkcal/hr
13.55
16.82
BFW supply flow
T/hr
12
15
Latent heat of vaporisation @ 98C,40 bar
Mkcal/hr
6.48
8.1
Heat carried by BFW
Mkcal/hr
1.176
1.47
Steam Produced
T/hr
11
13.65
Enthalpy of SH steam at 350C, 32.5 bar
kcal/kg
750
750
8.25
10.2375
Heat carried by SH steam
Mkcal/hr
Table 3. Comparison of parameters
8
7. CONCLUSION:
During high naphthenic feed processing in CCRU the delta T increases beyond 2500C resulting
higher OBSG inlet temperature above 8250C.During this BFW steam coil bypass further can be
raised beyond 3T/Hr. The above scheme is used in recovering heat from OBSG Flue gas
producing additional steam of 2 T/Hr resulting savings of ` 2 Cr/Annum.
8. LIST OF FIGURES:
S.No
Figure description
Page
1
Temperature profile of OBSG
5
2
OBSG original condition
6
3
OBSG current status
7
4
Temp profile of OBSG after scheme
8
9. LIST OF TABLES:
S.No
Table description
Page
1
2
3
Gas data
Steam data
Comparison of parameters
4
5
8
Microcrystalline waxes from Industrial polyolefinic by product
Manisha Sahai, Ajay Kumar and Sanat Kumar
CSIR-Indian Institute of Petroleum, Dehradun-248005
[email protected] ( Tel: 0135-2525794)
Microcrystalline waxes are an important class of petroleum products, which have some special
properties as compared to common (paraffin) wax, derived from petroleum, like high melting
point, high oil retaining tendency and good adhesiveness, as a result of which they find wide
applications in various industries like rubber industry, tyre industry, cable and cosmetics industry
etc. These are mixtures of hydrocarbons of 35 to 50 or higher carbons and having molecular
weight between 500 to greater than 700. Microcrystalline waxes are mainly obtained from
petroleum heavy distillates and tank bottom sludges. However due to surging demand of middle
distillates these heavy distillates fractions are now increasingly being used for the production of
middle distillates as a result of which the world market is in short supply. In India the entire
demand of microcrystalline waxes is around 12TMTPA, which are being met through import.
These waxes have a large potential market and a very strong competitive capability. So there is
considerable growing interest in finding alternate sources for these types of waxes that are cheap,
easily available, and environmentally friendly.
In the present study polyolefinic byproduct of HDPE production plant has been used as starting
material. These polyethylene byproduct waste has lower molecular weight as compared to high
density polyethylene, hence has relatively lower mechanical strength. Since these by products are
clean and homogeneous and also present in bulk at a single location hence these can be a useful
source for value added products. These by products have been depolymerized/degraded in an
inert atmosphere using additives to obtain waxes similar to petroleum derived microcrystalline
waxes. The conversion of these by products following chemical recycling results in the
formation of liquid, gaseous and solid hydrocarbons. The conversion was carried out in a
cylindrical glass reactor fitted with the condenser system to collect the vapors. Reaction was
carried out with 300 grams of feed and catalyst /additive concentration varying from (1- 10gm).
The reaction system was deoxygenated by flushing of nitrogen for 30 minutes. These have been
degraded at temperature 370–420°C for specific duration (3-8 hrs). The degraded product
obtained after the completion of reaction mainly consists of waxes, liquid and gaseous
hydrocarbons. The yield of wax lies between 72-89 weight % while the liquid yield varies
between 11-33 wt%.
The study indicates that industrial polyethylene by product can be a very useful source for
obtaining alternatives to petroleum derived waxes. Microcrystalline waxes meeting standard BIS
specification (Type D) have been obtained from these sources. Light paraffin oil obtained as by
product is a valuable product, which can be used as petroleum solvent.These industrial
polyethylene byproducts can be a very useful source for obtaining alternatives to petroleum
derived microcrystalline waxes.
1.0 INTRODUCTION
The plastics are preferred over conventional materials, because of their functional and economic
benefits to its end users (1). With the ever increasing demand, the disposal of waste plastics, both
pre consumer as well as post consumer waste is a big problem nowadays (2). The Pre-Consumer
Material is waste stream produced during a manufacturing process that has never reached the
end user and which is no longer being used for its intended purpose. These types of plastic waste
stream are homogenous, clean, and dry and also have high recycling rates. The preconsumer
waste stream can be categorized into categories viz, Post-Industrial waste and by product. The
post industrial waste is finished or partially finished product with no end use but has a recycled
content without any further processing. The by-Product (Co-Product) is the material that does not
meet technical specifications, and/or there is no market for the material. In this work the
Industrial polyolefin waste which is a byproduct produced from HDPE plant been used as the
starting material. There is a significant quantity of these wastes available which being clean as
well as devoid of various additives and hence provide an added advantage as no purification
steps are involved. So researchers worldwide are working to have a viable solution to this
problem. Therefore, new methods of recycling to convert plastic waste into useful products and
energy like chemicals, fuels, lubricating oil base stocks (LOBS) etc are being explored. One of
the viable product can be Microcrystalline waxes which are mostly imported and being solid
provides the ease of marketing.
Microcrystalline waxes are an important class of petroleum products, which have some special
properties as compared to common (paraffin) wax, derived from petroleum, like high melting
point, high oil retaining tendency , good adhesiveness ,and has a higher tensile strength as a
result of which they find wide applications in various industries like Cosmetics, Pharmaceuticals,
Protection of Plants, Fruits, Cheeses, and Vegetables; Food Packaging, Textiles, Paper,
Fiberboard, Wood, Potting Compounds for Condensers, Floor Polishes, Furniture, Skis, Leather,
Rust Prevention, Rubber, Printing Inks, and Lubricants etc. Microcrystalline Waxes are nontoxic and non-irritating to the skin and eyes. They are compatible with mineral waxes, vegetable
waxes, esters, and all of their oils. These are mixtures of hydrocarbons of 35 to 50 or higher
carbons and having molecular weight between 500 to greater than 700 (3). Microcrystalline
waxes are mainly obtained from petroleum heavy distillates and tank bottom sludges. However
due to surging demand of middle distillates these heavy distillates fractions are now increasingly
being used for the production of middle distillates as a result of which the world market is in
short supply. In India the entire demand of microcrystalline waxes is around 12TMTPA, which
are being met through import (4). These waxes have a large potential market and a very strong
competitive capability. So there is considerable growing interest in finding alternate sources for
these types of waxes that are cheap, easily available, and environmentally friendly.
In view of the above, therefore attempts have been made by us to convert the Industrial
byproduct of HDPE plant into microcrystalline wax range hydrocarbons by following additive
promoted degradation.
2.0 EXPERIMENTAL:
2.1 MATERIALS:
Pre consumer waste is from HDPE plant waste and production industry.
2.0 PROCEDURE:
CHARACTERIZATION OF FEED:
The feed containing no additive or filler and insoluble in all the recommended solvents at
ambient temperature was characterized for the following properties:
2.2.1 Differential Scanning Calorimetric (DSC) Analysis:
The thermal properties of feed were determined by using Perkin Elmer model Diamond DSC.
The measurements were carried out in aluminum crucible taking ~ 4mg sample with a heating
rate of 10oC/min. The data obtained on the stated parameters for feed has been given in table-1
Table-1
Properties
FEED
TM,OC
Peak
Area
HJ/g
90.55 109.02 116.75
94.74 112.28 117.97
3.659 60.09
2.496
-.962 16.3680 .6569
2.2.2 THERMO GRAVIMETRIC ANALYSIS:
The thermo gravimetric analysis (TGA) of feed was done on a Perkin Elmer make thermo
balance. For this about 10 mg of feed sample was taken and heated from 50 to 550 oC @ 10 oC/
min. The thermo gram between weight % & temperature and DTG curve of the same presented
in Figure -1
Figure-1 TGA& DTG thermo gram for Industrial polyolefins by product
The conversion of the feed was carried out in a 500 ml cylindrical glass reactor fitted with the
condenser system to collect the vapours .Reaction was carried out with feed (300gm) and
catalyst A concentration varies from (1- 10gm). The reaction system was deoxygenated by
flushing of nitrogen for 30 minutes. The waste plastic is heated at temperature 370–420°C for
specific time (3-8 hrs). The resulting products were liquid hydrocarbons, waxes and gases (5),
which were analyzed for different properties. The process schematic is shown in figure-2
2.2.3 CONGEALING POINT & NEEDLE PENETRATION:
These properties of feed were determined following standard procedures as per ASTM –
D 938-70(IP D76/70) and ASTM -D- 1321-76/IS 1448(P.93)1979 respectively. The melting
point was observed to be 100-110oC
2.3 PROCESS DESCRIPTION
There is no standard apparatus for studying the degradation of waste plastic and
researchers have used their own assembly. A prototype glass assembly was designed and got
fabricated at the laboratory. It contains a vertical type glass Reactor. This reactor is further
connected to three jacketed coiled condensers again through spherical joints. The fourth was an
ice trap which in series was connected to aqueous KOH/ NaOH solution. The condenser has
provision of receivers connected at their respective ends. The uncondensed gases have provision
to be measured through a wet gas meter. Finally the uncondensed gases can be flared or vent to
atmosphere. The temperature of the reactor is controlled by PID controller. The temperatures of
condensers are maintained by low temperature circulation bath.
Figure -2: Process Schematic for obtaining waxes
2.4 CHARCTERISATION TECHNIQUE FOR PRODUCTS
To evaluate the compositional and structural properties the samples have been analyzed using
FTIR, NMR, and GC etc. For quantitative information on the functional groups present in waxy
samples the samples were analysed using FTIR by making thin film on KBR pellets on Thermo
Nicolet-8700 spectrometer. An FTIR spectrum of a typical wax sample is shown in figure-3. The
wax samples subjecting to 1H NMR (Figure-4) were collected on a Bruker Avance III 500 MHz
NMR spectrometer operating at 500.13 MHz resonance frequencies for 1H, using 5 mm broad
band probe.To verify the carbon number distribution of the waxes , carbon number distribution
was done using G.C.
Figure-3 IR Spectra of typical wax
Figure-4 1H NMR spectra of typical wax
The waxes are characterized for their physicochemical properties like Drop melting point
Kinematic viscosity, Needle penetration, oil content, refractive index, ash content, UV
absorption etc. using standard methods. Drop melting was done using method D127/IP 133
The melting point is the temperature at which the first drops fall from the tip of thermometers
while it was heated. Kinematic Viscosity was measured using ASTM method D445, Viscosity of
majority of waxes varies from 3-6 mm2/s at 100oC.Congealing point was measured using
ASTM: D938/IP 76. The congealing points of waxes are invariably lower than their Drop
melting point. Needle Penetration was done using ASTM method D1321 .Oil content was done
using d721/ IP 15 a weighed amount of sample is dissolved in MEK and chilled to -31oC to
precipitate wax, the solvent oil mixture is filtered Refractive index was done using Instrument
by Anton Paar, refractive index is used for determination of carbon types composition .
2.5 DETERMINATION OF NATURE OF WAX
A scheme for classifying waxes as either paraffin, semi-microcrystalline, or microcrystalline is
based on the ASTM -TAPPI equation
n2D = 0.000194 3t + 1.3994
……………Equation-1
where t is the congealing point temperature in °F (ASTM D-938, IP 76). if the value calculated is
less than refractive index the nature of the wax is paraffinic in nature .while Semimicrocrystalline wax and microcrystalline wax are characterized by refractive indexes greater
than those given by the above equation and by viscosities at 210°F of less than 10cSt for semimicrocrystalline waxes or greater than 10cSt for microcrystalline waxes (figure-5).
1.4400
MICROCRYSTALLINE WAXES
B
Refractive Index at 212 o F
(Viscosity at 210 oF 10 cSt or higher)
AND
SEMI-MICROCRYSTALLINE WAXES
(Viscosity at 210 oF below 10 cSt)
1.4300
1.4200
A
1.4100
100
120
140
Congealing Point, o F
160
180
200
Figure-5 ASTM-TAPPI classification of petroleum wax
3.0 RESULT AND DISCUSSION
The DSC thermo gram of feed (Table-1) shows three transitions temperature respectively
at 90.55, 109.02 and at 116.75 oC. The transition at 109.02 involves maximum enthalpy & area
and is very close to the melting point (100-110oC) of the feed. This observation indicates that
feed sample is made up of at least three components which vary not only in their physical
properties but also in their structural properties.
The analysis of TGA and DTG curves (Figure-1) indicated that the degradation of feed
starts around 120 oC and the maximum loss in weight is observed between 300-480 oC. The DTA
curve also conforms to this deduction. These observations also conform to DSC deduction that
FEED contains at least three components varying in their physical properties. It can also be
deduced from TGA/DTG curves that the temperature from 450 to 500 oC would be most
favorable for carrying out degradation.
The degraded product after the reaction is mainly consist of waxes, liquid and gaseous
hydrocarbons. The table- 2 indicated the yield wt% of the waxes and liquid hydrocarbons
obtained after degrading feed under different experimental conditions. The yield wt% of wax lies
between 72-89%.while the liquid yield varies between 11-33 wt percent.
Table-2
Wax
W1
W2
W3
W4
W5
W6
W7
Yield of wax , Yield of lighter
wt%
hydrocarbons
85
15
89
11
83
17
79
21
78
22
79
21
67
33
Congealing
point of wax ,oC
89
85
87
82
79
83
72
Nature of wax
Microcrystalline
Semi-Microcrystalline
Microcrystalline
Semi Microcrystalline
Semi Microcrystalline
Semi Microcrystalline
Microcrystalline
3.1 STRUCTURAL PROPERTIES:
3.1.1 CARBON NUMBER DISTRIBUTION:
The carbon number distribution of the wax is in between C15 to C59
3.1.2 INFRARED SPECTROSCOPIC ANALYSIS:
The Spectra shows:




The C-H strong stretching band between 3000-2850cm-1, its scissoring between 14701450cm-1 ,its rock, methyl between 1370-1350cm-1 and its rock, methyl between 725720cm are all indicative of the presence of alkane
The strong stretch between 1642-1628 cm-1 and strong band between 1000-650 are
indicative of the presence of alkenes in all the samples. But they make only small % of
the total amount.
The finger print region has band at 1641(medium) -alkenes, 1467(strong) -alkanes,
1377(medium)-alkanes,991(medium)-alkenes,909(strong)-alkenes and 721(medium)phenyl ring substitution respectively are strong indication of the presence of higher
alkanes and alkenes in the samples.
It is most probable that N-H,O-H,X-H and C-=C are absent
So it can be inferred from the spectra the waxes are mainly consist of alkanes with a little
amount of unsaturation indicating the presence of alkenes. Whereas no hetro atom is
present in the product
3.1.3 1H AND 13CNMR:
The NMR spectra shows that the wax samples contain long chain of carbon component
and in them the CH2 chain is being the most predominant. Unsaturated protons (aliphatic
olefins) are less than 1%. The remainder of the content of waxes was composed of
aliphatic saturated hydrocarbons. No oxygenated species such as carboxylic acids,
aldehydes, ketones, ethers, or alcohols were detected by either 1 H or 13C NMR
spectroscopy The Compositional analysis by NMR spectroscopy shows the presence,
olefinic and paraffinic protons. Aromatic protons are absent. These have been confirmed
by 13 C NMR analyses also.
3.2 NATURE OF WAXES OBTAINED
Petroleum wax is generally of two types, the paraffin and microcrystalline which are obtained
from petroleum distillates and residue/sludge respectively. However since the nature cannot be
obtained directly, the waxes have been classifies using ASTM-TAPPI method. Three waxes have
been characterized for R.I and viscosity and their nature determined using equation-1 is shown in
table-3.
Table-3
S.No
Sample name
1
W1
Congealing Refractive Kinematic
point(oC)
Index( at
Viscosity
o
100 C)
(at 100oC)
83
1.434205 7.77
2
W2
89
1.434094
9.78
3
W3
87
1.433750
16.25
Nature of Wax
SemiMicrocrystalline
SemiMicrocrystalline
Microcrystalline
It is observed that the wax W1 and W2 are semi-microcrystalline in nature .The W3 is
Microcrystalline in nature.
3.3 COMPARISON WITH STANDARD WAXES
Waxes are graded according to their melting point (ASTM D-87, IP 55) and oil content (ASTM
D-721, IP 158).The melting point and oil content of waxes of waxes (ASTM D-87, IP 55) have
significance in most wax utilization. The wax has been compared with the standard BIS
specification. Wax have been characterized in detail and are shown in table-4 The Drop melting
point of the wax is 93oC which is in accordance with the Type -D as per BIS specification .
Congealing point of the microcrystalline waxes is around 87oC, the Microcrystalline waxes have
higher molecular weights, Microcrystalline waxes have greater affinities for oil than paraffin
waxes .
Wax shown in Table-4 meet the required BIS specification for Type D wax.
Table-4
S.No Properties
value
1
2
3
4
5
6
7
8
87%
93oc
87oc
0.003
Nil
Nil
1.5
1.0
Yield wt%
Drop Melting Point
Congealing Point
Ash wt%
Acidity
Saponification value
Oil content
Needle penetration
BIS
specs(TypeD)
85-95
0.03
nil
nil
1.5
1.0
4.0 CONCLUSION:
The study indicates that industrial waste plastics can be a very useful source for obtaining
alternatives to petroleum derived waxes. Microcrystalline waxes meeting standard BIS
specification (microcrystalline wax of type D) have been obtained from these sources. Light
paraffin oil obtained as by product is a valuable product, which can be used as petroleum solvent.
References:
(1) Plastic - Wikipedia, the free encyclopedia, https://en.wikipedia.org/wiki/Plastic
(2) Interpreting Pre-Consumer Recycled Content ClaimsPhilosophy and Guidance on
Environmental Claims for Pre-Consumer Recycled Materials.
www.recyclingstar.org/wp- ontent/.../08/Preconsumer_Recycled_content_claims.pdf.
(3) Paul Ratnasamy,K. S. Anand,D. C. Gupta; “Structure and properties of microcrystalline
waxes”, Journal of Chemical Technology and Biotechnology , Vol 23,(1973) Pages 183–
187.
(4) Hydrocarbon Vision 2030 for Northeast India, page157
( http://petroleum.nic.in/docs/visiondoc2030.pdf).
(5) D. P. Serrano, J. Aguado, and J. M. Escola ; “Developing Advanced Catalysts for the
Conversion of Polyolefinic Waste Plastics into Fuels and Chemicals”, ACS Catal., Vol 2
(2012) , 1924−1941
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
Lead –Lag arrangement in Cycle Water Treating Unit (DI-Unit) in MEG Plant
- Utsav Shankar
Syed Ashfaque Ali
Abstract
The cycle water treating unit, also known as De-Ionization (DI) unit in MEG plant is designed to remove
small amount of salts, acid, caustic, and CO2 from lean cycle water.
This treating unit consists of two Nos. of filter guards followed by two exchanger trains with a common degassifier. Each train comprises of One Cation exchanger and one Anion exchanger. The normal route of lean
cycle water, which is to be treated, is;
Filter Guard >>> Cation Exchanger >>> De-Gassifier >>> Anion Exchanger >>> Treated water tank.
One Train is in line while the other train is in idle condition for 41 hours after its regeneration of 7 hours as
the total cycle time is 48 hours.
We modified this route and put the 2nd train in series with the 1st one, leading to a cycle time of 60 hours,
i.e., increase of 12 hours of cycle time from the original design. This modification resulted in less
regeneration thus leading to a saving in utilities and chemicals required for each re-generation.
An estimated benefit of 2.7 million INR has been achieved per annum plus reduction in waste water
generation due to each re-generation of beds.
Page | 1
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
1.0 Introduction:
The Cycle Water Treating unit, also known as De-Ionization (DI) unit in MEG plant is designed to remove
small amount of salts, acid, caustic, and CO2 from lean cycle water.
Lean cycle water is the media in which ethylene oxide (EO), present in the cycle gas coming from EO
reactor(s) is absorbed along with the above mentioned salts, acids, etc. Only traces of CO2 get entrapped
in the cycle water. Remaining quantity of CO2 is absorbed in carbonate solution.
The DI unit consists of the following components.
(1) Micro filter section
(2) Cation exchanger with chemical dosing section
(3) Vacuum Degasifier with degasifier bottom pumps
(4) Anion exchanger with chemical dosing section
1.1 Micro Filter (G-541 A/B) -- 2 sets
Purpose: The purpose of this filter lies in removing suspended solids (SS) contained in cycle water in
order to prevent the ion exchange resin from being contaminated. The removal efficiency is 98% removal
of particle at > 10micron
Page | 2
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
1.2 Cation Exchanger (R-541 A/B) -- 2 sets
Purpose: This unit is aimed at removing sodium acetate of 463 ppm contained in cycle water
Process: Each of Cation exchanger contains wet strong acidic resin. The process stream is treated by
passing it through one Cation bed. The treated process stream is sent to the degasifier. The exhausted
Cation exchanger is switched every 48 hours and resin is regenerated by acid chemical. Regeneration is
accomplished in maximum 8 hours and normally 4 hours. Prior to regeneration, the process stream from
the spent bed is forward displaced to the process in series through the train of Beds (Cation and Anion)
that is being regenerated by means of DM water.
1.3 Vacuum Degasifier (C-541) -- 1 set (Common)
Purpose: This unit is aimed at reducing free CO2 of maximum 250 ppm contained in cycle water
1.4Anion Exchanger (R-542 A/B) -- 2 sets
Purpose: This unit is aimed at removing acetic acid of about 279 ppm contained in cycle water.
Process: Each anion exchanger contains wet weak basic resin. The process stream is treated by passing it
through one anion bed. The spent bed is switched every 48 hours and regenerated. The spent Anion Bed is
regenerated after the completion of the regenerations of the spent Cation beds. Regeneration is
accomplished in a maximum of 8 hours and typical regeneration time is 4 hours. Prior to regeneration, the
process stream from the spent bed is forward displaced to the process in series through the train of Beds
(Cation and Anion) that is being regenerated by means of DM water.
Each train of DI unit is put out of service based on operation time, DI outlet flow stream conductivity, pH,
UVs. Generally, DI unit performance remains such that all the analysed parameters remain within
specifications as given by the MEG plant Licensor. So, the maximum operation time of each train of DI unit
is 48 hours.
Inlet cycle water condition:
1) Flow rate 30㎥/hr
2) Temperature 37℃
3) Pressure 7.0 kg/㎠(g)
4) Containing ion
- Cations : 463 ppm wt as sodium acetate
- Anions : 279 ppm wt as acetic acid
Regeneration Steps:
1) Displacement
2) Backwash
3) Settle
4) Chemical Feed (H2SO4 for Cation bed and NaOH for Anion bed)
5) Slow Rinse
6) Fast Rinse
Page | 3
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
Route: The normal route of cycle water bleed is Micro filter >>> Cation exchanger >>> Degasifier >>>
Anion exchanger >>> Treated water tank.
Original Operation Philosophy: One train comprising of a Cation exchanger, common degasifier
and an Anion exchanger is in service for 48 Hrs., while the other train remains idle for 41 hrs., after
getting regenerated for 7 hrs.
Page | 4
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
The Modification
Objective: To increase operation time of the beds by operating both trains in series, thereby reducing
the number of regenerations of beds per annum which in turn results in reduction in chemical
consumption, DM water consumption and waste water generation.
New Operation Philosophy: Modifications were carried out by adding small segments to the existing
piping also new On/Off valves were procured and installed for catering to the new operation
philosophy. Some modification in the existing PLC was also done.
The DI unit is now operated with Lead-Lag philosophy, i.e., both the trains are in series and
simultaneous operation. All the analysed parameters, viz., operation time, DI outlet flow stream
conductivity, pH, UVs observed to be normal up to 60 hrs of DI unit lead lag operation. So the safe
operation time considered to be 60 hrs., and hence, increased from previously followed 48 hrs., which
in turn reduced the number of regenerations from 167/annum to 133/annum.
After every 60 hrs. of operation, the lead bed goes for regeneration and the lag bed comes and
occupies the lead position. After 7 hrs of regeneration, the fresh bed occupies the lag position.
During the above mentioned regeneration period of 7 hrs., the Lag bed, which has occupied the lead
position runs independently. Several lab tests were carried out before implementing this modification
to ascertain the quality of the water and no offset was encountered during this 7 hrs period.
Page | 5
Indian Oil Corporation Limited
Panipat Refinery and Petrochemical Complex
Modified Regeneration Steps:
1) Displacement
2) Backwash
3) Settle
4) Chemical Feed (H2SO4 for Cation bed and NaOH for Anion bed)
5) Soaking
6) Slow Rinse
7) Fast Rinse
Thus, major tangible benefits achieved are:







Reduction in number of regenerations
Reduction waste water generation
Reduction in use of DM water
Reduction in chemical consumption
Based on current utility and chemical price, the above benefits led to a saving of around 2.7 million
INR /annum
Major intangible benefit: Enhanced impurity removal capacity as more impurity is generated from a
High Selectivity Catalyst, presently being used in the plant.
st
This modification has been done for the 1 time in an MEG unit.
***
Page | 6
Raising the bar of TAN Limitation of LK Fraction for ATF Production
Biswajit Shown, Swapan Ghosh and Asit Kumar Das
Reliance Industries Ltd., Jamnagar, Gujarat, India-361142
Abstract
The conversion of light kerosene (LK) to aviation turbine fuel (ATF) requires removal of naphthenic
acids to protect the merox catalyst as well as to maintain the specification of ATF. The electro
coalescer prewash (ECP) used as neutralisation process is unable to handle some high acidic LK
streams as light deviation of the LK TAN can become cause for formation of tight emulsion in the
ECP, poisoning of downstream sweeting merox catalyst, ETP upset and finally off spec. production of
ATF.
Strongly basic anion exchange resin based heterogeneous neutralization process is able to avoid
emulsion formation due to contact of naphthenic acids from LK with dilute caustic solution as happens
in the case of ECP. The newly developed process is able to handle any range of acidity of the LK
streams to make it suitable for downstream sweetening merox process.
Ambient temperature operation, recycle of used chemical and easy regeneration, feed acidity level
flexibility and ability to avoidance of issues of existing procedure make the resin based acid
neutralization method a unique, efficient and cost effective. This also helps to widen the crude
processing window by blending more opportunity crude oils with high acidic LK fractions.
Introduction
Naphthenic acids (measures as total acid number, TAN) of light kerosene feed plays an important role
in the performance of merox catalyst in fixed bed sweeting process as well as product specification of
aviation turbine fuel (ATF). Naphthenic acids are removed as sodium naphthenate by treating with 2-3
wt. % caustic solution in the prewash electrostatic coalescer section before entering to the merox
reactor. A typical neutralization reaction of naphthenic acid with caustic is shown below.
R – (CH2)n – COOH + NaOH
R – (CH2)n – COONa + H2O
Generally the prewasher works well with feed TAN 0.05 mg KOH/gm (max.) and TAN level of 0.015 0.02 mg/gm is achieved at the outlet. But during processing of light kerosene with TAN > 0.2 mg
KOH/gm, coalescer outlet TAN exceeds min. requirement and affects merox reactor catalyst
performance in the downstream. Another important limitation of processing of High TAN light
kerosene feed is formation of stable rag layer in the interface, which is required to be drained to OWS
and results in severe problem in ETP.
In this article an efficient and economical alternate technology, based on heterogeneous acidity
neutralization process has been described for trouble free processing of LK stream with very high TAN
(0.4 instead of designed 0.05 mg KOH/gm). This heterogeneous base is capable to avoid formation of
rag layer formation and water carry over to downstream Kero Merox process unlike conventional
electro coalescer process using caustic solution.
Resin Based Acidity Reduction Process
Generally, basic anionic resin act as a heterogeneous base for acidity neutralization. Naphthenic acids
present in crude oils are ionized as naphthenates (RCOO− carboxylic acid anion) and H+. Exchange of
said carboxylic acid anion from the mobile phase is occurred with OH− ion present in the stationary
phase (heterogeneous base).
In this process highly cross-linked strongly basic macro porous anion exchange resin was used for
heterogeneous acidity neutralization. The counter ion OH− is replaced by naphthenic acid anion
(RCOO−) during passing of naphthenic acid containing LK feed through the bed. The exhausted resin
bed was regenerated again by passing mild alkali solution at ambient temperature.
Experiment
Experimental studies were conducted for acidity reduction of LK fraction using 150 ml. anion
exchange resin bed with bed diameter 2.5 cm. for determination of resin regeneration cycle, utilization
of active sites and optimization of alkali strength. Acidic LK feed was passed through the activated
resin bed at 0.5 to 3 hr-1 LHSV depending upon the feed acidity. At breakthrough the exhausted resin
bed was regenerated by alkali treatment.
Results and Discussion
Figures -1 & 2 show the study result. The acidity of the feed LK fraction was 0.5 mg KOH/gm. Resin
bed was considered as exhausted when the treated LK TAN was more than 0.02 mg KOH/gm. Overall
96% acid reduction was observed. In seven runs average break through point was 86% run length of 1st
run and average vol. of treated LK was ~98 times of resin bed volume. This implies excellent acid
removal capacity of resin bed as well as consistent regeneration efficiency.
Figure-1: Run length % w.r.t Run-1
Figure-2: Vol. (times of bed vol.) of treated
120
100
80
Vol. of Treated LK
(Times of Resin Bed)
% Run Length w.r.t 1st Run
100
60
40
20
80
60
40
20
0
0
Run-1 Run-2 Run-3 Run-4 Run-5 Run-6 Run-7
Run-1 Run-2 Run-3 Run-4 Run-5 Run-6 Run-7
Techno-Economical Advantages
The above mentioned process has several technical advantages over the conventional process as
follows:
 An alternate continuous process for reduction of acid content of hydrocarbon feed.
 Is able to neutralize any range of acidity unlike conventional electrocoalescer process.
 The resin bed is regenerable and can be used several years.
 Heterogeneous neutralization process helps to avoid formation of tight emulsion, which eventually
avoids the frequent upset of effluent treatment plant (ETP).
 Cost effective due to ambient temperature and atmospheric pressure operation.
Conclusion
The resin based heterogeneous acid neutralization process is highly efficient and cost effective due to
ambient temperature-pressure operation and ability to avoidance of adverse effect of electro
coalescence prewash process which is being practiced currently. This method is also capable to handle
any range of acidity of the light kerosene (LK) feed at the cost of bed regeneration cycle length. Simple
regeneration process also takes care opex and decrease ETP load by recycling alkali solution as well as
wash water. Moreover this method will allow refiners to process more opportunity crude oils which
contain high TAN at light kerosene fraction.
A process for the production of Low Poly-Aromatic Hydrocarbon (PAH)
Rubber Process Oil (RPO) from Distillate Aromatic Extract (DAE) and other
blending streams – Value upgradation
V.Selvavathi, M.Lavanya and R.Krishnamurthy
Chennai Petroleum Coporation Limited, Manali, Chennai-600068.
Introduction :
The Rubber Process oils are essential components used in the manufacture of tyres
and allied industries. They are generally added to natural or synthetic rubber to
enhance uniform dispersion of the fillers and to improve the flow characteristics of
the rubber. They contribute directly in stabilizing the quality of the tyre and
improving road safety. Thus RPO is an essential component for the technical
performance of the tyre and in particular for its road adherence (or grip) properties.
The Rubber Process Oils constitute about 5 -12% of the total weight of the tyre.
Applications of RPO:


They are used as carrier oils, plasticizers, dust control agents and processing
aids.
Apart from the rubber industry, other industries which make use of process
oils are plastics, printing inks, pharmaceuticals, food, cosmetics, chemicals,
textiles, gas processing etc.,
Advantages of adding RPO:




Decreases the mixing temperature during processing
Prevents scorching or burning of the rubber polymer
Reduces the power consumption and mixing time
Helps in improving uniform dispersion of ingredients and better extrusion due
to its lubricating property
Classification of RPO :
Various types of RPO are used depending on the type of rubber and the end product
application. They are largely classified as (i) Aromatic oils (ii) Naphthenic Oils
and (iii) Paraffinic Oils
Aromatic Oils: The aromatic oils are normally prepared from aromatic extracts
obtained in the refinery stream.
They are suitably blended to meet stringent
specifications in the manufacturing of tyres, tubes, hoses, calendar sheets etc..
Page 1 of 6
Naphthenic Oils: Naphthenic oils are characterized by their unique colour stability,
solubility and good thermal stability and ideal for moulded articles, slippers, LPG
tubes, floor tiles, etc..
Paraffinic Oils: These are oils containing predominantly branched and linear
paraffins with various viscosities and are used extensively in ethylene-propylene
rubbers.
Global Demand for Rubber Process Oils:
The global RPO market in 2013 was close to 3 million tons which may likely to
increase up to 3.5 Million tons by 2017. The aromatic oil accounts for 57% of the
total demand (aromatics comprising DAE, TDAE, RAE, TRAE, and MES ) where as
the
naphthenic and paraffinic RPOs contributed to 29% and 14% of the total
demand respectively. Aromatics are the most compatible RPOs for tyre industry due
to which its consumption share is high. The global RPO market is expected to grow
at a CAGR (Compound Annual Growth Rate) of around 2.5% upto year 2023, and in
Asia it is expected to grow at a CAGR of 3.7%, the highest in the world. The Indian
market for RPO is fourth largest in the world next only to US, China and Japan.and is
expected at CAGR of 2.5%
Regulatory Issues:
European Union (EU) regulating authority specified from January 2010 onwards that
DAE should meet
Registration, Evaluation, Authorisation and restriction of
Chemicals (REACH) specifications in the manufacturing of tyres. According to
REACH specifications, DAE should have less than 10 ppm of eight numbers of
identified PAH compounds while Benzo(a)pyrene is restricted to 1 ppm. This
regulation virtually eliminated the usage of DAE in the EU and adversely affected
demand in other countries that export tyre products to the EU.
Indian Tyre Sector:
India has emerged as global hub for automotive industry. To support and sustain the
significant growth in automobile sector around the globe, all the leading tyre
manufacturers are increasing their capacities to fulfill demands of Indian and
overseas markets.
In order to fulfill the European market demands, Indian tyre manufacturers are
exploring various alternatives for substituting DAE with Low PAH RPO.
Low PAH Rubber Process Oils
Typically an aromatic oil contains at least 70%wt. of total aromatics and from 10
to 15 wt% of Poly Aromatic Hydrocarbons (PAH). Although these oils show good
compatibility with SBR and other rubbers, they have the disadvantage for being toxic
and carcinogenic, largely due to the high PAH content.
Page 2 of 6
Hence, a need arises for creating a new substitute which is of low toxic & noncarcinogenic in nature. The new substitute low PAH RPO can be used as a process
oil for rubber compounds in general and more particularly, aromatic rubber
compounds.
The heavy DAE stream from lube refinery cannot be used directly as low PAH RPO.
The Rubber Process Oil meeting low PAH RPO grade can be produced from heavy
DAE by (i) distillation followed by blending with other streams available in Refinery
and by (ii) Re extraction of DAE.
Low PAH RPO:
Low PAH RPO grade is equivalent to Export grade RPO which has a stringent
specifications for PAH (Polycyclic Aromatic Hydrocarbons). According to this spec.,
the product should have less than 10 ppm of Benzo(a)pyrene, Benzo(e)pyrene,
Benzo(a)anthracene, Chrysene, Benzo(b)fluoranthene, Benzo(j)fluoranthene,
Benzo(k)fluoranthene, Dibenzo(a,h)anthracene while Benzo(a)pyrene restricted to 1
ppm.
We have followed two routes for achieving the above specifications :
1. By adjusting the IBP of the distillation to remove the 8 listed PAH components
from heavy DAE (BN extract) and blending with rich paraffin / naphthene
streams.viz: UCO / Foots oil
2. By re-extracting the Heavy DAE (BN extract)
Route 1 : Distillation and blending
The boiling point of the 8 PAH compounds are falling in the range of 438-524°C and
hence heavy DAE (BN extract) stream was distilled off upto 540°C to remove the 8
listed PAH compounds using ROFA TBP-Potstill unit. To meet the other required
specifications like viscosity, density, aniline point, RI, VGC etc., 540°C + stream of
BN extract is blended with either Foots oil or UCO in a specified ratio. The blended
streams are characterised for their properties and found to meet the required
specifications.
(a) 90 wt.% of 540°C + stream of BN extract + 10 wt.% of Foots Oil
(b) 91 wt.% 540°C + stream of BN extract + 9 wt.% of UCO
Export grade RPO spec and the results obtained for our blends are given in Table-1.
Page 3 of 6
Table-1 Low PAH RPO
Description
Export grade Low
PAH
RPO
RPO Spec.
(Blend of 540°C+ BN
extract and Foots
oil)
Appearance @ 15°C
Kin. Viscosity @ 100°C, cSt
Refractive Index @ 20°C
Density @ 15°C, gm/cc
Aniline Point, °C
Pour point, °C
Flash Point, °C
V.G.C @ 100°C
Sulphur content, wt%
Water content, wt%
PAH profile by HPLC
Benz (a) pyrene, ppm
Sum of the 8 identified PAHs,
ppm
Low PAH RPO
(Blend of 540°C+
BN extract and
UCO)
Dark liquid
Dark liquid
Dark liquid
25-75
1.545 – 1.565
0.96 – 1.0
<75
30 max
250 min
0.8 – 0.91
4 max
-
73.6
1.5599
0.9895
67
21
>250
0.9005
3.46
<0.1
72.69
1.5548
0.9890
70
24
>250
0.8992
3.43
<0.1
1 max
10 max
<1
<10
<1
<10
Route 2 – Reextraction of heavy DAE
Heavy DAE viz., BN extract was re-extracted using NMP. Extraction studies were
carried out using Single Stage Batch LLE system with the following conditions.
Feed
Solvent
S/F Ratio
Temperature
Extraction time
Settling time
: BN Extract
: NMP
: 2.0
: 90 °C
: 1 Hr
: 1Hr 30 Min
After extraction, the raffinate stream was collected. The solvent free raffinate was
characterised for its properties and the results are found to meet the Export grade
RPO (Table-2).
Page 4 of 6
Table-2 Low PAH RPO
Description
Appearance @ 15°C
Kin. Viscosity @ 100°C, cSt
Refractive Index @ 20°C
Density @ 15°C, gm/cc
Aniline Point, °C
Pour point, °C
Flash Point, °C
V.G.C @ 100°C
Sulphur content, wt%
Water content, wt%
PAH profile by HPLC
Benz (a) pyrene, ppm
Sum of the 8 identified PAHs, ppm
Export
Spec.
grade
RPO Low PAH RPO (Reextracted
BN
extract)
Dark liquid
Dark liquid
25-75
1.545 – 1.565
0.96 – 1.0
<75
30 max
250 min
0.8 – 0.91
4 max
-
77.3
1.5548
0.969
69
24
>250
0.884
3.78
<0.1
1 max
10 max
<1
<10
Low PAH RPO meeting Michelin Raw material Specifications :
Another Environment friendly grade Low PAH RPO meeting the specification
requirements of Michelin can be met by blending the Lube Oil Base stocks (LOBS)
available in the Refinery. The details are given in Table-3
Table-3
Characteristics
Specific gravity @ 15°C
Viscosity @ 100°C, cSt
Viscosity @ 40°C, cSt
Aniline Point, °C
Flash Point, °C
Pour point, °C
%Ca
DMSO Ext. (IP 346), wt. %
Michelin rawmaterial spec.
0.895-0.925
13-17
150-200
85 to 100
≥ 220
≥0
11-17
< 3.0
Hydrofinished
LOBS
0.9001
14.5
151.59
101
255
-3
16.9
0.59
Value Addition :
The current RTP prices of the Blending Components and the market price of Low
PAH RPO are given in Table-4. As seen from the prices, there is a good value
addition in the production of Low PAH Rubber Process Oil.
Page 5 of 6
Table - 4: Prices of Blending streams and formulated Low PAH RPO
Cost of Blending streams (Approximate)
RTP Prices as on
Blending streams
Foots oil
UCO
BN Extract
Landed cost of Low
PAH RPO if Foots oil
is used as blending
component
Landed cost of Low
PAH RPO if UCO is
used as blending
component
Hydrofinished LOBS
Cost in ` per Kg
30.05
32.74
35.40
*38.36
*38.68
50-55
Landed cost of Low
PAH RPO by
blending
hydrofinished LOBS

Cost of Low PAH RPO
(Approximate –
Exclusive of taxes)
Cost in ` per Kg
100
100
54.25
Inclusive of 10% operating cost
Summary :

The Low PAH RPO is Environmental Friendly which meets the stringent
REACH specifications of Europe.

Low PAH RPO equivalent to “Export grade Rubber Process Oil” can be
produced by the following two routes :
 Adjusting IBP of BN extract to remove the 8 listed PAH compounds
and blending with either Foots oil or UCO to meet the other properties
like viscosity, aniline point, RI etc.,
 Re-extracting BN extract to remove the 8 listed PAH compounds

Michelin raw material specifications can be met by blending hydrofinished
LOBS

Low PAH RPO has high value addition and the price is 2 – 2.5 times the
selling price of the refinery streams used for blending
Page 6 of 6
Improvements carried out in Oil Movement and Storage section
to augment capacity of HPCL Mumbai Refinery
A.B.Chattopadhyay, Ashok Kumar, Ganeshraj G, Syed Arif Hussain*
HPCL Mumbai refinery
1. Introduction.
The primary role of oil movement and storage section (OM&S) is to receive
crude/products from designated agencies/units through the tankers and pipe lines, store
them in the respective tanks and transfer the same to designated agencies (terminals /
customers) or refinery process units. In Mumbai refinery, crude is received through
tankers and pipelines. Products are evacuated out through pipelines. The refinery shares
crude and some product pipe lines with other industry members like port, TATA power,
BPCL, AEGIS after paying applicable wharfage charges. Hence, timely evacuation of
products helps in optimizing the costs and avoiding inconvenience to other industry
players. OM&S also maintains continuous and close coordination between
marketing/refinery planning for the shipment/receipt of the products to maximize asset
utilization. Various losses like ocean loss/ transfer loss/ processing loss are monitored
and minimized to achieve the fuel and loss target of refinery. Treatment of effluents is
also undertaken.
This paper deals with details of development of process specifications, detail engineering,
& operating philosophy of all the initiatives implemented in OM&S section which
contributed immensely in achieving the above stated objectives. HPCL Mumbai refinery
is in the process of expanding the refining capacity from 7.5 to 9.5 MMTPA. In this
project, it is being ensured that all the improvements described are implemented in
project stage itself.
2.
Crude tank modifications.
a) Migration issues
Crude oil storage tanks receive PG Crude through marine tankers. This crude oil is
normally processed in FR Combination unit. Single isolation valve was provided in
recirculation discharge, crude suction and upper and lower suction lines. Normally, the
recirculation discharge is lined up to any of the other crude tanks for emergency in crude
unit. Also, some time is spent in closing the isolation valves. Due to passing of the single
isolation valve, crude migrates from one tank to the other or to the ship.
To overcome this issue, it was proposed to provide additional block valve. But due to
unavailability of space and to reduce manual operation, DBBV with MOV is being
implemented for positive isolation and faster operation from centrally controlled OM&S
DCS.
b) Single deck roof undulation issue
Floating roof tanks (capacity of around 72 TMT each) were of single deck roof type with
pontoons on periphery and buoys welded. Due to the large diameter of the tanks, there
used to be undulations on the roof, thereby creating an uneven stress on the weld joints of
the single deck. Thus, there was a chance of crude coming out of the rooftop creating an
unsafe situation. Risk is increased during monsoons.
Single deck roof have been converted to double deck roofs with compartmentalization to
avoid any leak and roof submergence.
c) Roof drain facility
Due to high rainfall in the region, the reliability of the roof hose drain is very critical.
These crude tanks were having swivel joint type roof hose drain pipes, which fail if the
joints are not properly fixed. If the drain fails, the oil starts coming out of the roof drain
along with the rainwater, leading to unsafe situation and OWS over filling.
Existing swivel type roof hose drain pipe are being replaced with armor clad flexible roof
hose drain pipe for better reliability as per section C.3.8 of API650 guidelines or
unbonded flexible hose with inner carcass as per API 17J suitable for crude oil
environment.
d) Crude mixers
Crude layering and sludge deposition is another issue frequently encountered in crude
tanks. Even though mixer provision was available, reliability of the mixer with fixed
shaft was low, leading to frequent failures and outage of the tank to completely take out
the mixer for maintenance and repairs, leading to ullage and economic concerns.
The fixed shaft Butterworth old mixers are being replaced by latest swivel joint type
impeller mixers, as per ASME B 16.47.
3.
Low discharge rate ex crude tanker.
Mumbai refinery has a dedicated pipeline from port manifold to 6 numbers of crude tanks
for receiving the crude oil. The 36” crude receipt line gets connected to 24” BH crude
receipt line inside individual tank dyke and joins tanks in a 24” nozzle. It has been
observed that maximum crude receiving rate from marine tanker is 5,500 kls/hr.
Suction spool piece and nozzle size in the tank are being increased to 36” to increase the
flow rate from 5500 kls/hr to 7000kls/hr.
4. Floating roof primary seal.
Floating roof tanks were provided with foam seal liquid mounted/ vapour mounted
system. In this, compressible foam is held in an elastomeric envelope, which traps the
vapor at the top of the rim gap or prevents evaporation by resting on the stored product.
While regular movement of the tank roof, the elastomeric envelope has a tendency to leak
and lead to fugitive emission from tank.
As per the statutory guideline requirement of fugitive emission gazette of India
notification (Part2, Section 3i), all the foam seals in floating roof tanks are being changed
over to mechanical shoe seal. The benefits are
• Elimination of seal gap due to levelness of mechanical shoe.
• Maintaining uniform pressure of shoe on tank shell.
• Compatibility with all products resulting in uniformity.
• Sealing arrangement doesn’t cause friction while rubbing against the shell.
5. Conversion of black oil jetty line.
In FCC units, VGO from the primary processing units is cracked to produce valuable
products like LPG, naphtha, diesel blending streams. To utilize the excess capacity
available in FCCs, VGO needs to be imported. As VGO being a high pour material, it is
not possible to receive the same through the existing submersible black oil line available
from Jawahar Dweep. There was an existing 24˝ Tata Power import pipeline from New
Pir Pau Jetty up to Tata Power. This line is further connected with 14˝ and 10˝ LSHS
lines of HPCL, which was earlier used to supply LSHS from HPCL-MR to Tata Power.
During a crude unit revamp shutdown in the past, VGO was brought from Vizag using
this line. However, due to restriction of line size at HPCL, the discharge rate was not
more than 400 TPH and hence the parcel size had to be restricted to 5 TMT only. Tanker
discharge rate is around 1000 TPH. This pipe line was not available for importing high
pour material due to safety issues.
A new 20˝ line up was laid and extended in its first part from Y Junction bridge to old Pir
Pau Jetty (approx. 1 km) and further to new Pir Pau Jetty (first Chemical Berth). Second
part (approx. 2 km) was laid within the compound of the Mumbai Port Trust (MbPT).
Further extension of the line up to second chemical Berth is current under construction by
MbPT and also resizing of the existing 14˝ line from MbPT up to MR to 20˝ line is being
taken up in a subsequent phase. In order to strengthen the logistic infrastructure, this new
line up is proposed for using as multi-product line for handling various products like
VGO, LSHS and FO. Processing of various streams imported from this new line is
estimated to give a net margin benefit of about Rs.39.0 Cr per annum.
6. MUPL, LPG Import and storage at M/s Aegis logistics.
HPCL & BPCL have jointly constructed and commissioned an LPG pipeline of 0.8
MMTPA capacity from Mumbai to Uran (MUPL). The projected production of LPG at
MR is 35-40 TMT per month, out of which around 15 TMT per month is used at Mahul
BDU for bottling & around 10-15 TMT will be later used by MR for production of
propylene post PRU setup after MREP. Thus for transfer to Uran from Mumbai refinery
10 TMT per month would be available against 40 TMT per month capacity of the
pipeline. To meet the LPG demand of the region, HPCL is importing around 30 TMT per
month at the terminal of M/s Aegis Logistics Ltd. & around 10 TMT per month at JNPT
Uran terminal of BPCL. Further, there is no source of LPG production coming in near
future in the region. Also, post commissioning of MUPL, no tanker truck is allowed
inside Mumbai city as per govt directive HPCL will have import LPG at Uran at higher
charge of 1,505 INR per MT compared to terminal charges of 719 INR per MT at Aegis
Logistics terminal.
It was decided to continue the usage of the import facilities of Aegis Logistics terminal
and transfer of products from their terminal to HPCL refinery mounded storage bullets by
laying 8” LPG pipeline from Aegis terminal to Mumbai Refinery to meet the full capacity
of MUPL of 40 TMT per month. Under synergy of Industries, M/s Aegis Logistics Ltd
agreed to lay the pipeline Benefits to HPCL are
• Tariff at terminal would Rs 719 per MT (1,505 per Mt at BPC Uran). Thus there
will be a saving of Rs 28 crores per annum for a volume of 30 TMT per month.
• Full capacity utilization of MUPL.
• No dependency on Uran terminal facility.
• Product will be made available directly from Mumbai refinery to Mahul BDU
during shut down/shortage in production of Mumbai refinery. Presently, such
requirement is to be moved by tanker truck at a cost of Rs 1,100 per MT.
7. Multi product piggable line.
At present there are 3 white oil lines (of 14”, 24” and 12” dia size), a 20” dia black Oil
line, a 36” dia crude receipt line and a 14” water flushing line along with a 14” LSHS line
going to MbPT Pirpau manifold. From there onwards, lineups are available for Jawahar
dweep Island, OPP Jetty, NPP Jetty, Marketing terminals and Aegis Logistics. The three
white oil pipeline and 20” black Oil pipelines pass through a residential area outside the
boundary of Mumbai refinery raising safety and pilferage concerns. The lines also cross
over rail tracks, due to which the maintenance activities get delayed as it require
permissions from Rail authorities. As HT electricity cables of TATA power also passes
through nearby area, there is a concern of static charge in the atmosphere. Apart from
these bottlenecks, line decongestion need to be seen for future MREP developments.
Feasibility study has been carried out to convert multiple White oil (3 no’s of 12”, 14”
and 24”) / Black Oil line of 20” into two multiproduct piggable pipeline of 24” dia from
Mumbai Refinery OM&S to MbPT manifold boundary limits. It will bypass the
residential reducing the chances of pilferages and enhancing safety of populated areas.
The HT lines and the Railway line corridor will be bypassed in the lineup routing of the
piggable pipelines enhancing the safety of the transfer pipelines to jetty. This project will
lead to decongestion and debottlenecking helping in refinery expansion project. Detailed
engineering study is in progress.
8. Water draw off facility
Before feeding to CDU for processing, the received crude is given settling time and the
settled water is drained out through the water draw off (WDO) line of the tank. This
water may contain some crude oil or emulsion, if adequate settling time is not given due
to urgency or ullage. Hence water draw off of PG and BH Crude is routed to baby tanks
for further separation where again water is taken out after giving settling time. Similarly,
water has to be drained from the ATF product storage tanks to prevent water carryover to
ASF. Water from WDO goes to baby tank for further settling and separation.
The water draining was done manually. The operator carried out water draw off until
emulsion starts to appear and completes the operation. During the Water draw off
operation, some quantum (approx. 10%) of oil used to get entrained along with the water
drained and go to separator for oil recovery. In this process oil is lost due to evaporation.
Automatic water draw off valves have been installed at the existing water draw off tanks.
These work on the principle of difference in specific gravity of oil and water. They close
automatically after sensing the presence of product emulsion leading to lower losses
during WDO operation, reduced opex of ETP and slop reprocessing. As the performance
of these valves was as excellent, they are now being installed in all product tanks.
9. Connection of FR and FRE FBS.
Due to increased demand of SCN product and its better realization as compared to
Naphtha and HSD, it was decided to maximize SCN production. Earlier, low Sulphur
FRE FBS was routed to Tk 370/312 via SCN r/d or to HSD back blending via HSD E3
r/d. Tk 370 has got provision of routing to the customer. FRE FBS production is 24 TMT
per month and FR FBS production is 11 TMT per month. Requirement in HSD is approx.
14 TMT per month and hence balance FBS needed to be exported.
Post revamp of crude unit, FR FBS was also meeting SCN specs. During HSD 4
production, FR FBS could not be routed to HSD due to sulfur specifications and hence,
FR FBS had to be routed to SCN r/d. Hence, connection of both FRE FBS and FR FBS
was made in Offsite section in order to facilitate both FBS to meet market demand.
Benefit realized is approx. 2.43 Cr per annum.
10.
Propane tanker truck loading facility.
Current production of Propane from HPCL-MR is 660 MT per month. Propane is
received in bullets from propane (PRU) unit. The total capacity for the two bullets is
500MT. During crude unit or PRU shutdowns, LPG production gets affected and PRU
feed is reduced, leading to lower propane production. Low propane inventory leads to
reduction in feed to lube units.
Hence, during PRU shutdown, propane tanker truck was procured from outside, and
propane was unloaded from them to the bullets with specially made unloading facility,
sustaining the lube unit PDA feed. Estimated benefit was approx. Rs 75 Lakhs.
11. Flare purge control and gas recovery
A new control system was built to control the fuel gas flow as purge and back burn purge
for flare stability as well as for eliminating smoke during normal scenario as well as
emergencies like unit trip and hydrogen flaring.
New flare gas recovery compressors have been installed in place of older compressors.
Benefits are:
•
•
•
Higher handling capacity of 3384 m3/hr with a Molecular weight of 10.3
compared to older 4480 kg/hr with Molecular weight of 45.78.
Improved reliability.
Flare loss has been reduced from 0.3 wt% to 0.1 wt%.
12. Conclusion.
HPCL Mumbai refinery has taken various initiatives to implement the state of art
technologies for loss and opex minimization and enhancing refinery throughput. Fuel and
loss has been brought down by over 0.5 wt%. Mumbai refinery is also tracking various
statutory guidelines being published and is ensuring compliance of the guidelines.
21st Refining Technology Meet, 2017
Analysis of complex flows in a petroleum refinery
Muffazal Badshahwala, Rahul C Patil, Ajay Gupta, Asit Das
Abstract
Conventionally, the flow behaviour in a vessel is described by ideal behaviour, completely
mixed or plug flow. Though the concept of ideal flow can provide key information about
reactor/ unit operation performance, the flow patterns in a vessel depends on its geometry, and
may deviate significantly from the assumption of uniform ideal flow. Without an accurate
understanding of the complex flow patterns present within a given system, designers have to
rely on conservative assumptions to ensure that equipment meet design requirements.
Computational fluid dynamic (CFD) is often used to study the non-ideal flows inside the
equipment and has been successfully used in many applications to improve performance of the
operations either at the design stage or during on-stream operation.
The petroleum refinery and petrochemical plants have been linked with inherent complications
with respect to complex flows involving multiple phases at various unit operations, reactors,
pipelines etc. It is thus often necessary to study these complex flows to improve unit’s
reliability to reduce downtime and subsequently improve overall margin. This paper presents
couple of case studies which have been implemented in petroleum refinery and exploit some
of the flow models offered by commercial CFD software.
Case study 1.
Improve the heat transfer in Air preheater (APH) unit.
The heat transfer enhancement plays a key role in improving energy efficiency. Non-uniform
flow of flue gas exchanging heat with process air flowing through tubes in an air preheater
(APH) leads to inefficient heat transfer and results in low outlet temperature of process air.
Eventually, the fuel fired in the furnace increases and the furnace efficiency reduces. The
objective of this study is evaluate the existing air duct design for apour ng ution of air
(non-ideal flow) and subsequently to suggest internals in the air duct which would significantly
reduce the vapour ng ution in the vapour flow to APH.
The geometry considered for simulations are shown in figure 1 and 2. The steady state pressure
based solver was used to solve the flow equations. The turbulence in the flow was modelled
using standard k-ε model. The simulations were continued until the residue fell below 10-5
indicating convergence of the flow problem.
Figure 1: Case 1; Base case (APH unit without vanes)
Figure 2: Case 2; Modified design (APH with vanes)
The standard deviation of vapor mass flow at the plane just below air preheater tubes were
computed for two cases and considered as the measure of uniformity of flow distribution.
Lower the standard deviation, more uniform (ideal) is the flow. The plane was divided into 16
sections to estimate the standard deviation in mass flow.
Results and observation:
As mentioned earlier, the vapor mass flow for each case is quantified in terms of standard
deviation. It was seen that case 1 have more variation of mass flow around the average value
over case 2. The normalized standard deviation values for each case are given in table 1.
Table 1: Standard deviation in Vapor flow distribution
Mass flow
Case 1
Case 2
Standard Deviation
0.52 X Base
0.08 X Base
Min mass flow
0.05 X Base
0.85 X Base
Max. mass flow
1.72 X Base
1.10 X Base
Average mass flow
Base
The vapor flow can also be visualized through the contours across the plane z=0. Contours for
each cases are shown in fig 3-4.
Figure 3: Relative velocity for case 1
Figure 4: Relative velocity for case 2
Conclusions:
Standard deviation of around 8% in mass flow as compared to around 52% in base case was
observed) without causing much significant pressure drop. Thus it can be concluded from the
above analysis that introducing vanes assists in attaining uniform distribution at the inlet of air
preheater.
Case Study 2
Determination of liquid entrainment in an inlet entry devices for a flash zone of a
distillation column.
The function of inlet entry devices in a flash zone of a distillation column is to separate apour
and heavier liquid component from two phase flow of gas and liquid and distribute the apour
uniformly into the column. The non-uniform distribution (non-ideal) of apour will cause
localized higher velocities leading to carry over of heavier component of liquid to the column
bed/tray just above the entry device which subsequently deteriorates the quality of lighter
product. It is thus necessary to ensure the uniform distribution of apour from the entry device
to ensure lower entrainment. Two entry device designs namely, tangential entry apour horn
(see fig 5a) and radial entry apour horn (see figure 5b) have been compared for degree of
entrainment of liquid and uniformity of apour flow.
(a)
(b)
Figure 5: (a) Tangential entry apour horn (TEVH) and (b) Radial entry apour horn (REVH)
Since, the liquid fraction is very low (less than 1 wt%), the Eulerian-Lagrangian approach is
utilized to study the flow of apour and liquid through entry devices. Thus initially the apour
flow through the devices was established. The standard deviation of apour velocity was
compared at plane just above the entry device and the plane just below the column bed to
analyse the apour flow uniformity. The particles were then tracked in the domain using
Eulerian-Lagrangian approach. The particles are assumed to be moving with the fluid and oneway coupling is enabled in the simulation. The path of the particle which hits the walls of flash
zone entry device gets truncated and the remaining particles escape either from the bottom
chimney tray or outlet at the top. The injected inert particles have uniform droplet diameter.
The study has been conducted for different diameters varying from 10 to 200 µm. Coalescence
and droplet breakup in domain or after hitting the wall surface is not considered in the
apour ng and simulation. The total percentage of number of particle streams truncated and
escaped from the bottom chimney tray is evaluated for each devices. The steady state pressure
based solver is used to solve the flow equations. The apour flow is assumed to be ideal gas.
The turbulence in the flow is modelled using realizable k-ε model.
Results and discussion:
The standard deviation over average velocity in y-direction for each cross section is considered
as a criterion for uniformity of apour distribution. Lower the standard deviation over the
average velocity in particular cross section, uniform is the apour distribution. The standard
deviations over average velocity in y-direction in two planes for two entry devices, TEVH and
REVH are shown in table 2. The velocity profiles at these two planes are shown in figure 6 and
7 respectively.
Plane
TEVH
REVH
Plane just above entry device
1.8 x Avg. Velocity
1.6 x Avg. Velocity
Plane just below the column bed
0.2 x Avg. Velocity
0.1 x Avg. Velocity
(a)
(b)
Fig 6: y-Velocity profile at plane just above entry device (a) REVH (b) TEVH
(a)
(b)
Fig 7: y-Velocity profile at plane just below column bed (a) REVH (b) TEVH
It can be seen from table and figures that the flow profile in REVH is better as compared to
TEVH, owing to flow path in respective devices. In REVH flow is concentrated at the centre
of the column as against the TEVH where flow is concentrated at only once section of the
column where flow restriction is present. To study the effect of components other than yvelocity also, on the droplet trajectory the inert particles are injected from the inlet surface and
their trajectories are tracked in Lagrangian frame of reference. The number of particle streams
injected is equal to the number of meshes on the inlet surface. The plot of percentage of particle
streams trapped and escaped from the bottom of entry devices are shown in figure 8.
TEVH
REVH
100
% Particles Captured
90
80
70
60
50
40
30
0
25
50
75
100
125
150
175
200
Droplet Size (µm)
Fig 8: Percentage of particle streams trapped and escaped from the bottom of entry devices
Greater is the percentages of particles hitting the walls of flash zone entry devices, better are
the chances liquid droplets to coalesce and settle at the bottom of flash zone entry device.
REVH showed 100% particle collection efficiency at droplet size of 100 micron and above. It
is clear that TEVH is less efficient in de-entraining the liquid droplets in two devices considered
in the study.
Conclusions:
It is shown that REVH is better device with respect to vapor distribution as compared to TEVH.
The particle tracking study indicated 100 % particle collection efficiency for REVH at 100
micron droplet size and above. Results for TEVH indicate poor performance as compared to
REVH for all droplet sizes. Overall, REVH is better than TEVH in distributing vapor uniformly
and de-entraining liquid droplets.
Motive Steam pressure optimization via Vacuum Ejector Modelling
Authors
Vaskar De
Mayur Talati*
Mayur Tikmani
Reliance Industries Limited,
Jamnagar Mfg. Division (Gujarat, India)
www.ril.com
* Lead Author; e-mail : [email protected]; Ph : 08511124205
Abstract
Multistage vacuum system is common part of any Vacuum distillation unit (VDU). Each ejector
is designed to pull VDU overhead suction vapours using motive steam (at given pressure)
using critical flow motive steam nozzles. This design provides an opportunity to optimise
steam consumption via manipulation of the motive steam pressure.
The governing limit for the motive steam pressure (and thus steam flow) reduction is the
location of sonic shock wave (or sonic boost, pressure front across which the fluid velocity
redu es elo Ma h . i the diffuser’s throat. For the su essful operatio of eje tors,
sonic shock wave must reside within the throat (i.e. straight pipe portion of the diffuser). This
can be ensured by maintaining the ejector back-pressure below allowable Max. Discharge
Pressure (MDP).
In the present work, a mathematical model is developed from the first principle to predict
MDP for the given set of the ejector(s) considering the ejector geometry, operating condition
and estimating efficiency of various nozzles at varied motive steam pressure. The predictions
of the programme has then been successfully implemented in an operating vacuum
distillation unit to optimise the motive steam consumption. Also from the rigorous model,
simple correlations (suitable for DCS calculations) are developed to automate the
opti izatio approa h y pla i g a al ulatio lo k i DCS a d pro idi g advised oti e
steam set-pressure to the operator for real-time optimization.
The implementation of this approach has resulted into the reduction of about 20% motive
stea i a three stage eje tor syste at RIL’s Va uu Distillatio U it. I pri iple, the
approach is applicable to any ejectors (sonic motive nozzle design) and thus has high
replication potential across refineries.
Working Principle of the Steam Ejector:
Ejectors are characterized by their ejector curves. Typically this consists of two curves. One is
the suction curve and the other is the maximum discharge pressure (MDP) curve. Figure 1
shows a typical ejector curve.
Suction curve gives the relationship between suction
pressure and the suction load. It is independent of the
discharge pressure as long as the back pressure is
below the MDP curve. MDP curve gives the maximum
allowable back pressure at which the ejector operation
can be kept stable for a particular suction load. From
Fig. 1, it can be seen that both suction load and MDP
increases with suction pressure increases.
Fig. 1 Performance Curve of an ejector
In an ejector, power comes from steam
pressure and flow. Since the steam nozzle is a
critical flow orifice, flow is set by supply
pressure. The higher the steam pressure, more
the steam flow rate and greater the energy for
compression. Figure 2 shows the effect of
changing motive steam pressure on the ejector
performance curve. It can be seen that, as the
motive steam pressure is increased, the MDP
curve shifts upwards and vice versa. Increase in
motive steam increases the steam consumption.
Fig. 2 MDP variation of an ejector with motive
steam pressure
Problem Analysis:
The above theoretical analysis shows that MDP is a function of suction pressure and motive
steam pressure. The relationship is:
a) Lowering suction pressure lowers the MDP at given motive steam pressure
b) Lowering motive steam pressure lowers the MDP at given suction pressure
As discussed earlier, for a stable ejector operation, back pressure should be lower than MDP.
Thus a reliable / dependable estimate of MDP is a must for ejector motive steam
optimization. This is because vacuum break is not desirable due to operational difficulties and
VGO loss. In current operation, motive steam pressure is only slightly decreased below design
conditions. This leads to sub-optimum operation of steam ejectors.
This study is aimed at developing an ejector model based on thermodynamics principles that
can be used to evaluate and optimize the performance of steam jet ejectors.
Model Development
In order to arrive/establish at a dependable and simpler correlation between the motive
steam pressure and Max. Discharge Pressure (MDP), complete ejector model is developed
from the fundamentals, the co-efficient are regressed from the available performance
curve, actual ejector performance data and relevant literature. The reasoned assumptions
are appropriately made in the operating range of our interest, to keep the mathematical
treatment as simpler and as linear as possible for ease of implementation. This section
provides an account of the considerations went into the said modelling, resulting into
required relation.
Use of constant pressure mixing paradigm has been chosen for modelling mixing of the
motive steam and entrained vapour due to its favourable match with the experimental data
as cited in several published literature [1] [2] [4]. The pressure and velocity profile for this
model is as shown in Figure 3.
Fig. 3 Variation in stream pressure and velocity as a function of location along the ejector
The major assumptions and equations of the model are:
1. The motive steam (primary fluid, p) expands isentropically in the nozzle with efficiency
ηn.
and the vacuum column overhead vapour (entrained fluid, e) expands isentropically in
the suction chamber.
Where M is the Mach number, P is the pressure, γ is the isentropic expansion coefficient,
2. The mixing process is modelled by one-dimensional continuity, momentum and energy
equations assuming steady state conditions.
Where w is the entrainment ratio and M* is the ratio between the local fluid velocity
to the velocity of sound at critical conditions.
3. The constant pressure assumption implies that the pressure between points 2 and 4
remains constant. Therefore, the following equality constraint applies:
4. Pressure increase across the shock wave at point 4 is given by
5. Pressure lift in the diffuser
Where, ηd is the diffusor efficiency
6. Friction losses are defined in terms of the isentropic efficiencies in the nozzle (ηn),
mixing chamber (ηm), and diffuser (ηd)
7. Nozzle efficiency (ηn) only varies with motive steam pressure and is independent of
entrained fluid suction pressure (Pe) with
ηn = 1 (at design motive steam pressure)
8. The mixing chamber efficiency is inversely related to the projected nozzle exit area
ηm ∝ 1/A2
Where projected nozzle exit area (A2) refers to the area formed at 2 after mixing of
motive steam and entrained fluid.
9. The diffusor efficiency and nozzle projected area is a function of suction pressure
ηd, A2 = f(Pe), at given motive steam pressure
10. Overall adiabatic efficiency (ηo) varies with suction pressure and motive steam pressure
ηo =f ( Pe, Pp)
Model Inputs
a) Development of MDP curve at design motive steam pressure
In order to obtain the parameters of the above developed model at the design motive steam pressure,
available performance curve of ejectors were used.
Available Data points were fitted in to equation of the form:
---------------------eq (1)
Non-linear regression was used to obtain parameters a, b, c, & d. Figure 4 shows comparison of the
actual data (from performance curve/field observation) and the graph predicted by above equation
for ejector
(Design)
Fig. 4. MDP curve of J01 ejectors
b) Estimation of diffusor efficiency and nozzle exit Area/pressure at design motive steam
pressure
It has been assumed that diffusor efficiency (ηd) and projected (effective) nozzle exit area (A2)
varies with entrained fluid suction pressure (Pe) at a given motive steam pressure. This is
because, with the change in suction pressure, the mixing pattern and the mixing energy will
change. It is expected that the increase in suction pressure will create more jet effect at
nozzle exit plane and thus decreasing effective nozzle area i.e. A2. Also increase in suction
pressure should increase mixing in the diffusor section and thus increasing ηd.
As regards to motive nozzle efficiency (ηn), it is reasonable to assume that ηn remains constant
with change in suction pressure as long as motive steam pressure is kept constant. This is
because, efficiency of motive nozzle should depend on conditions within the nozzle and
therefore should remain constant at fixed motive steam pressure. In the nozzle exit plane,
sonic region exists and therefore, projected nozzle exit pressure varies inversely with effective
nozzle exit area.
To estimate ηd, P2=f(Pe) , we used the MDP curve developed by eq (1) and the suction pressure
s apa ity data eje tor’s perfor a e ur e to tu e our ther ody a i odel, y aryi g
ηd, P2. This yielded an equation of the form y=mx + c between (ηd, Pe) and (P2, Pe)
c) Estimation of ejector efficiencies at different motive steam pressure
Estimations for overall/motive nozzle efficiency at different motive steam pressure were
made by literature [3] references and plant operating data.
Overall efficiency graph
1
0.95
0.9
0.85
Overallefficiency
0.8 (lit)
0.75
0.7
Range of Interest
0.65
0.6
2
4
6
8
10
12
14
16
Motive steam pressure (kg/cm2g)
Fig. 5 Variation of overall efficiency of a typical ejectors with motive steam pressure
Results
The above discussed correlations were used in the model to predict the MDP curve at
different motive steam pressure. Simple polynomial correlation was developed for predicting
the required motive steam pressure from ejector suction pressure and ejector back pressure
keeping a fixed margin from MDP. These correlation was implemented in DCS to
automatically set the ejector steam pressure by measuring the plant conditions.
Motive Steam Pressure
P1
P2
P3
P4
P5
Fig. 6 J01 MDP curve at different motive steam pressure
Optimization of the ejector and Benefit evaluation
As shown earlier, with the model, we have developed MDP curve of the ejector at different
suction pressure and different motive steam pressure. With this MDP curve available, we
can lower the MP steam pressure to a point where MDP is higher than back pressure by a
pre-set margin.
In order to demonstrate savings potential of the proposed philosophy, we applied this
model to optimize MP steam consumption in one of our crude unit’s ejectors for the time
period of t o years i.e. Ja ’ to De ’ o a o thly a erage asis.
Following graph depicts the total MP steam savings cumulative for the ejector system (time
period : Ja ’ – De ’ .
Base Case
~ 20%
Proposed Case
savings
Fig. 7 MP steam consumption in base and proposed case
Conclusion
By theoretical modelling of ejector system a steam saving potential of about 20% in this
system was revealed. Relation of MDP vs Suction pressure and Motive steam pressure was
found out using this first principal model. The relation was condensed into a polynomial form
for easy implementation in DCS. It was demonstrated in plant that this change will lead to the
expected saving in operating cost.
References
[1] S. Ghorbanian and S. J. Nejad, Ejector Modeling and Examining of Possibility of Replacing
Liquid Vacuum Pump in Vacuum Production Systems, International Journal of Chemical
Engineering and Applications, Vol. 2, No. 2, April 2011
[2] R.B. Power, Steam Jet Ejector for Process Industries, McGraw-Hill, 1994
[3] F. Shinskey, Ch.4, Compressor control systems, Energy Conservation Through Control.
[4] H. El-Dessouky, H. Ettouney, I. Alatiqi, G. Al-Nuwaibit, Evaluation of steam jet ejectors,
Chemical Engineering and Processing 41, 2002
Application and Benefits of CombustionONE
Solution to Fired Heaters
A White Paper
TI 53A90A01-01E-A
Application and Benefits of CombustionONE Solution to Fired Heaters
Application and Benefits of CombustionONE Solution to Fired Heaters
Introduction
Sections
the introduction of the TDLS technology. The
CombustionONE solution unites a TDLS analyzer
Introduction
1
Managing Combustion in Natural Draft Fired Heaters
2
How It Works
5
CASE STUDY: A Major Gulf Coast Refinery
The Recommended Best Industry Practice
6
7
While fired heaters are used throughout refining and
with a dedicated control system and a safety system
petrochemical processes as the source of process
certified to meet FM NFPA and SIL 3 standards.
heat, they carry inherent risks and costs that make
operating without current technologies problematic.
The intrinsic value gained by adopting this new
To address safety concerns, Industry standards
combustion technology can be summarized by the
are upgrading their recommended practice for
following:
Instrumentation, Control, and Protective Systems for
Fired Heaters and Steam Generators. While not yet
Best Industry Practices
required by regulation, plants not meeting the best
Increased safety
Industry Practice guideline, will be at added risk in
The Technology
8
The Installation Process
9
the event of an incident on a fired heater. Since many
natural draft fired heaters do not meet this guideline
Longer fired heater life
with existing instrumentation and control systems,
an upgraded system consisting of recommend
technologies will be needed. Further, since most
natural draft fired heaters have only automated
Figures
Improved thermal efficiency
Lower greenhouse gas and pollutants emissions
Increased throughput
control of the fuel supply, not air, excess air is often
applied to the combustion process, reducing thermal
Figure 1: Diagram of a Typical Fired Heater
2
Figure 2: Features and Benefits of TDLS Technology
3
efficiency. According to ARC INSIGHTS, “Second
only to raw materials costs, energy is the leading
cost pressure currently affecting manufacturers.
New analysis techniques, such as Tunable Diode
Figure 3: TDLS Technology for Combustion Safety and Optimization
4
Figure 4: Basic Functional Capabilities of CombustionONE
5
Figure 5: Refiners Participating in the API 556 Task Team
7
Figure 6: CombustionONE - An Integrated Solution
8
Laser Spectroscopy (TDLS), can improve efficiency,
maximize throughput, reduce emissions, and improve
safety and reduce energy in combustion processes.”
Second only to raw materials costs, energy is
the leading cost pressure currently affecting
manufacturers. New analysis techniques, such
as Tunable Diode Laser Spectroscopy (TDLS),
can improve efficiency, maximize throughput,
reduce emissions, and improve safety and
reduce energy in combustion processes.
ARC INSIGHTS, INSIGHT# 2009-50MP,
November 2009
This paper highlights and describes the application
of a new combustion solution system by
Yokogawa called CombustionONE. By incorporating
Tunable Diode Laser Spectroscopy (TDLS)
technology, this new system has the ability to
simultaneously control air and fuel supply to fired
heaters by measuring average gas concentrations
across the high temperature radiant section.
Measurement of O2, CO and CH4 at high temperatures
and by cross sectional averaging using the gas
spectrum has only recently become reliable with
X
1
Application and Benefits of CombustionONE Solution to Fired Heaters
Managing Combustion in
Natural Draft Fired Heaters
The air supply for most fired heaters is natural
draft – not forced air - and these heaters normally
lack the extent of automation as the other process
units in the plant. Natural draft fired heaters, as the
name implies, use flue gas buoyancy to support
combustion. These heaters can be either cylindrical
or box type (see Figure 1 below). The buoyancy of
the flue gas (combustion product) relative to the
surrounding air is determined by the product of the
average density of the flue gas and the height of the
heater. Furnaces are designed to run at a pressure
of -0.05” to -0.1” WC at the top of the radiant
Application and Benefits of CombustionONE Solution to Fired Heaters
the fired heater, reducing its thermal efficiency. The
and CO at the top of the radiant section, ideally one
The low level of control on most fired heaters
lack of effective instrumentation to continuously
foot below the roof tubes. This is where combustion
is due, at least in part, to the historical lack of
and rapidly measure O 2 and CO in the combustion
reaction is expected to complete under all heater
reliable, effective instrumentation and automation
chamber of fired heater introduces considerable
operating loadings. This is an ideal application of
technology to simultaneously measure and control
safety risk. Apart from the simultaneous control of
CombustionONE, which employs new TDLS tech-
the fuel, gas concentrations, and the air/fuel ratio.
fuel and air concentrations, it is possible for fuel-
nology. Using the TDLS in concert with a dedi-
An oxygen (O 2) sensor is typically required at the
rich conditions to arise which increases the potential
cated controller, a cross section average O 2 and
stack base for thermal efficiency calculations, which
explosion risk. Note that under fuel-rich conditions,
CO concentration can be measured to determine
require total excess air. While operators attempt to
temperature/fuel controllers no longer work properly.
the right air/fuel ratio, rather than a localized spot
maintain ‘excess’ O 2 in the furnace for safety, the
amount indicated from an existing sensor may be
incorrect due to tramp air. In fact, it is possible that
the burners may be starving for air, despite excess
oxygen at the stack base. Because of the lack of
Apart from the simultaneous control of
fuel and air concentrations, it is possible
for fuel-rich conditions to arise which
increases the potential explosion risk.
air control, operators typically allow excess air into
section, whether a heater is natural draft, induced
draft or forced draft. Figure 1 provides a simplified
diagram of a typical fired heater.
measurement of these gases. Using an average O 2
Figure 1: Diagram of a Typical Fired Heater
and CO concentration produces safer burner control
and greater overall heater efficiency. Figure 2 lists
the features and benefits of TDLS technology, which
has distinct advantages over single point, in-situ
analyzers that may give false readings because of
The detection of combustibles - primarily CH4 - in
varying gas concentrations at different locations in
the radiant section of the fired heater is recom-
the fired heater.
mended by the American Petroleum Institute per API
Figure 2: Features and Benefits of
TDLS Technology
556; however, traditional analyzer technology cannot
be installed in the radiant section due to the high
Stack
“Inefficient combustion can be attributed
to the air/feul ratio. Too much excess air
(air rich) results in loss of efficiency and
increased NOx emissions, while too little
air (fuel rich) is downright dangerous.
Carbon monoxide measurements provide
an indication of fuel-rich conditions, while
oxygen measurements indiate air-rich
conditions. The optimal control point is the
lowest possible excess air value that does
not cause the system to enter an unsafe
condition or violate emissions limits.”
Feature
Benefit
measurements of CH4, O 2 and CO concentrations,
In-situ analysis
Sample conditioning not required
operators tend to allow excess air than necessary in
Fast response
Real-time data for control
Tunable laser
Interference-free analysis
temperatures. As noted above, without accurate
Damper
the heater, reducing its thermal efficiency.
Breeching
Convection
section
Shield section
Process Fluid In
Process Fluid Out
To properly control the combustion air,
CH4 and CO must be measured at the
top of the radiant section where combustion is completed, regardless of the
burner loading.
Non-contact sensor Suitable for operations in harsh environments
Optical sensor
Low maintenance
Source: ARC INSIGHTS, INSIGHT# 2009-50MP, November 2009
Fired heaters have two principal unsafe operating
conditions that must be avoided:
To properly control the combustion air, CH4 and CO
Coil
Radiant
section
Because of the lack of
air control, operators
typically allow excess
air into the fired heater,
reducing its thermal
efficiency.
must be measured at the top of the radiant section
•
Fuel Rich - where air is reduced, CO will be
where combustion is completed, regardless of the
produced by the burners and excess O 2 will be
burner loading. Note that O 2 and CO will coexist
lower, which results in excess fuel (CH4) from
within the flames, where the temperature may be as
the burners
high as 2,200 degrees F. Low NOx burners may use
delayed completion of combustion through staged
Burner
Natural Air Draft
•
Flame Out – where loss of flame results in rapid
air/fuel mixing, or external recirculation of cooler
loss of gas temperature. O 2 levels are high as
flue gas with combustion air, reducing peak flame
burner air is not reduced by combustion, and
temperature. In either case, to effectively control the
un-combusted fuel (CH4) is present
combustion process, it is essential to measure O 2
2
3
Application and Benefits of CombustionONE Solution to Fired Heaters
Application and Benefits of CombustionONE Solution to Fired Heaters
Consequently, continually measuring percent O 2 is
of CombustionONE TM, the embedded control and
critical to improving heater efficiency and maintain-
safety systems will ensure that these conditions are
ing safe operating conditions. When firebox con-
avoided or that combustion is extinguished and fuel
Capable of measuring the average gas concentra-
below illustrates the key components comprising the
ditions are unacceptable, i.e. high levels of CO or
flow is interrupted automatically if these conditions
tion across the radiant zone of the fired heater,
solution. Because the TDLS is a non contacting
combustibles exist, the effective combustion man-
are detected. The TDLS analyzer technology will
CombustionONE addresses both of the above less-
measurement - never touching the flue gas - and
agement solution must rapidly detect the condition
reliably respond to all “O 2 events”, where conven-
than-optimum conditions by simultaneously control-
has no moving parts, the whole system enjoys very
and initiate the appropriate response. In the case
tional sensor technology will miss most such events.
ling the fuel and air (O 2) supply based on 5 second
high reliability. The TDLS analyzer technology within
sample intervals. Measuring the gas concentration
CombustionONE has been operational on furnaces
in the radiant zone is also a requirement of API 556.
since 2003 without incident and most of those units
Figure 3: TDLS Technology for Combustion
Safety and Optimization
By optimizing air
flow control, O2
concentration is
typically reduced
from 6% to 2%,
increasing thermal
efficiency of the
furnace
How It Works
heater steam purge valves and fuel trip valves when
necessary. The simplified architecture diagram
Figure 4: Basic Functional Capabilities of CombustionONETM
Refinery Heaters
CombustionONE
Cool temperatures but CO reaction
is nearing completion
Concentration
Variations
Safety
System
Convection Tubing
TDL Receiver
TDL Transmitter
Radiant Tubes
(Radiant Zone)
Fuel System
Control Element
Burners
Gas Concentration
Measurement
(TDLS)
Field
Instrument
Control Systems
After CombustionONE
Before CombustionONE
4
Control Element
•
Higher costs as operators increase O2 flow to
avoid a fuel rich atmosphere
•
Reduced O2 and lower operating costs as the
fuel-air mixture is controlled
•
Unexpected demand for fuel, leading to unsafe
combustion conditions
•
Fuel is limited to the available air to prevent
unsafe fuel rich combustion
•
Greater risk during a process upset
•
Process upsets are handled with controlled
combustion conditions
•
The amount of excess air and unburned fuel may
not be assessed correctly during process upsets
or strong winds
•
O2 gas and unburned fuel concentration is
detected readily and controlled at optimum level
during upset conditions and strong winds
•
Wet steam introduced on start up, requiring a
steam purge, risking ignition failure
•
•
Shorter life of convection section with afterburning due to presence of combustibles
•
Measurements from the system include CO, CH4, O 2
have not required calibration. The analyzer has
and temperature. Using an average gas concentra-
full diagnostic capability, and if there is an issue,
tion produces safer burner control and greater over-
the analyzer will alert the operator. Also, the
all heater efficiency. By optimizing air flow control,
measurement signals from the TDLS are unaf-
O 2 concentration is typically reduced from 6% to
fected by the presence of other gases in the flue
2%, increasing thermal efficiency of the furnace.
Enforced drain removal from purge steam prevents unsafe ignition attempt
Shortening of tube life is prevented by maximizing radiant heat absorption and eliminating afterburning caused by combustibles
The CombustionONE system manages fuel flow and
arch draft through the existing plant DCS via Modbus, and combustion airflow directly.
The safety system with CombustionONE receives
Continuous measurements
of percent O2 is critical to
improving heater efficiency
and maintaining safe
operating conditions.
inputs directly from the furnace and operates the
5
Application and Benefits of CombustionONE Solution to Fired Heaters
gas, unlike sensor based technology. TDLS uses
gas analyzer to verify the results. Data collected
a path average measurement, as opposed to the
during this testing will be incorporated into the
traditional point measurement, making the value of
future safety shutdown system. The modular proce-
concentration much more accurate.
dural automation capability in CombustionONE will
CASE STUDY: A Major Gulf
Coast Refinery
Application and Benefits of CombustionONE Solution to Fired Heaters
The Recommended Best Industry
Practice
The most critical times of heater operation are at
start-up and shutdown. Recognizing that a much
faster, more reliable analyzer is required to measure
O2, CO and CH4 concentrations, a major Gulf Coast
refining operation has adopted the CombustionONE
solution for combustion management. Since
CombustionONE measures gas concentration in
the radiant section of the fired heater, the system
is expected to improve heater safety and overall
operational efficiency.
enforce safe operating conditions during start-up
and shutdown. Once the Safety Interlock System is
lished by the American Petroleum Institute, is a rec-
in place, CombustionONE will be able to detect and
ommended practice for “Instrumentation, Control,
interdict any unsafe operating conditions.
and Protective Systems for Gas Fired Heaters”. This
Natural draft heaters lack the capability to use
steam generators in petroleum refinery, hydrocarbon-
air to purge the heater. Instead, steam is used. If
processing, petrochemical, and chemical plants. API
the steam is not dry water will accumulate on the
556 specifically states the following regarding the use
burners/igniters, preventing ignition. The start-up
of TDLS technology:
sequence, part of the CombustionONE solution, will
Laser based technology for combustion
control (oxygen trim to air or air/fuel ratio
controller) is a design consideration for
heater applications where a single sample
point will not provide a representative
sample. It has a response time of ≤ 5
seconds and can measure across a radiant section up to 98.4 ft. (30 m). It is not an
ignition source to flue gas and requires no
reference air.
purge the condensate from the steam line, thus providing dry steam to purge the heater before ignition.
Results
CombustionONE has been operational since June,
2010 and the TDLS analyzers continue to operate
reliably with no maintenance needed. The operators
Solution
CombustionONE was installed with two Tunable
Diode Laser Spectroscopy (TDLS) analyzers for
measurement of O2, CO and CH4 concentrations
in the radiant section of the heater. A dedicated
CombustionONE controller uses these
measurements while feeding these values to an
existing DCS for monitoring and future control of the
heaters. The dedicated control hardware is equipped
to receive additional signals from the heater, which
will be used to control the airflow in the burners in
a subsequent phase of the project. Space has also
been allotted for a future safety shutdown system.
have been able to reduce the percent O2 by 1% to
1.5%, thus making the heater more efficient. The
upset conditions and the response of combustion gases will be analyzed to confirm the desired
response to unsafe conditions. This data can then
be correlated to the output from the existing stack
Task Team continues to refine the API 556 guideline.
API 556
vided representatives to the API 556 Task Team.
Figure 5: Refiners Participating in the API 556
Task Team
Company
minimum excess air. The TDLS measurements have
Not covered in this recommended practice are the
been verified by the existing stack gas analyzers,
following:
but with a percent O2 reading of 1% to 1.5% lower
than the stack gas analysis because the measure-
•
Oil fired and combination fired heaters
ments are taken in the radiant section. Furnace con-
•
Water tube boilers which consist of single or
ditions can now be controlled (or shutdown) quicker
multiple burners and are designed for utility
since CombustionONE is taking concentration
operation or where the primary purpose is steam
measurements at five second intervals in the radiant
generation (Covered by NFPA 85)
section. If there were to be an excess concentra-
•
Fired steam generators used to recover heat
Contribution
BP
• Chairman SCOICS
• Heater Design
Chevron
•
•
•
•
•
Protective System
Control, protective
General Instrumentation
Process Control
Heater Design
CITGO
•
•
•
•
Valves
Analyzers
Protective System
Instrumentation
and Analytical
ConocoPhillips
•
•
•
•
•
•
Analyzers
Design
Burners
Control
Safety
Valves
ExxonMobil
• General Instrumentation
• Process Control
Marathon Ashland
Petroleum
• Chairman SCOICS
• Combustion/general
• General
Instrumentation
Valero Corp.
• Analyzers
• Controls
furnace is now near optimum operating point – using
from combustion turbines (HRSG)
tion of CO or CH4 in the furnace, these gases can
be detected earlier versus the conventional stack
The CombustionONE system will be tested during
of the refining and petrochemicals industry, the API
Figure 6 lists the refining companies who have pro-
Over two years in development, API 556, to be pub-
guideline specifically applies to gas fired heaters and
Challenges
Composed of representatives from a cross-section
•
Oven and furnaces used for the primary purpose
gas analyzers, enabling the heater to be shutdown
of incineration, oxidation, reduction or destruc-
sooner and avoiding unsafe conditions.
tion of the process medium (covered by NFPA
86).
•
Water bath or oil bath indirect fired heaters (Covered by API 12K)
•
CO boilers, pyrolysis furnaces, e.g. ethylene and
hydrogen reformers, and other specialty heaters
6
7
Application and Benefits of Combustion Management to Fired Heaters
The Technology
•
A Combustion Interface Unit (CIU) consisting
of a cabinet and PC/Monitor to interface with
The CombustionONE Total Control solution is
Combustion Control and Safety System.
an integrated, self-contained system that can be
rapidly installed on any fired heater. Five principal
Two TDLS systems (transmitters and receivers)
components that comprise the solution are shown in
are typically installed in the radiant section of the
Figure 6 and described as follows:
fired heater, which is the most accurate location for
optimum combustion management. At minimum,
•
TDLS technology for gas concentration
one unit measures O 2 and the other CO and CH4.
measurements on 5 second intervals
Since fired heaters have differing configurations,
capacities, environmental and process conditions,
•
A dedicated Combustion Control system for control
custom mounting brackets are built to hold and
of fuel and air flows based on a fired heater model
position the laser module and detector module
across the radiant section. This is the best location
•
•
An OSHA compliant and SIL 3 Safety System to
to obtain the most accurate gas concentration and
prevent unsafe conditions from persisting
temperature measurements.
Sensing and actuation for additional
The dedicated combustion controller is used to si-
measurements and air flow control as needed
multaneously control fuel and air flows based on five
Figure 6: CombustionONE - An Integrated Solution
TDLS for Gas Concentration
Class 1, Division 2
Rated CIU Enclosure
Application and Benefits of CombustionONE Solution to Fired Heaters
second sampling measurements of average gas con-
•
centrations across the radiant section from the TDLS.
The specific fired heaters on-site where CombustionONE can be of highest value
Embedded with the controller are proprietary fired
heater combustion strategies that direct fuel flow at
•
the valve and air flow through the burner registers.
An OSHA compliant safety system is required on a
A note of the changes needed to each fired
heater that require detailed engineering
•
fired heater to meet regulatory requirements. The
The number of TDLS analyzers and their installation location on the target heater
safety system will safely shutdown the fired heater
if the main controller fails or if unsafe conditions
•
persist in the fired heater. The embedded safety
Order for installation of the CombustionONE
solution
system in CombustionONE is certified to meet
the requirements of NFPA, FM, ISA S84.01 and
•
A general project time-line and budget
IEC61508 and SIL3.
Since the CombustionONE solution is completely
Smart (self diagnosing) multi-variable transmitters
self-contained and requires little integration with
and valve positioners provide reliable and accurate
existing control systems and instrumentation,
performance and feedback of the combustion man-
the installation phase is typically straightforward.
agement solution. Additional SIL 2 transmitters are
Where mechanical modifications and/or upgrades
sometimes added on the fired heater to provide air
to the fired heater are necessary, Yokogawa acts as
flow measurements at multiple locations, ensuring
the prime contractor, partnering with engineering
safety and optimum performance.
companies and furnace manufacturers to provide
a single source for the project. The installation
Dedicated Controller
A Combustion Interface Unit (CIU) consisting of a
and commissioning processes occurs in 8 to 12
Class 1 Division 2 cabinet and PC/Monitor is an im-
weeks, depending on the size and complexity
portant component of CombustionONE solution. It is
of the fired heater. Installations on multiple fired
equipped with a 17” touch screen LCD for viewing
heaters on a site or across a fleet can also be
gas concentrations, long term analyzer trend data
easily accommodated. A systematic installation
and Standard Operation Procedures (SOP) for safer
plan that schedules the installation across multiple
heater startup and shutdown. The CIU unit interfac-
fired heaters offers greater economy of scale, while
es with Combustion Control and Safety System.
reducing the time to realize the benefits from the
CombustionONE solution.
The Installation Process
Sensing and Actuation
As a best practice, Yokogawa recommends
Safety Interlock System
commencing the installation of a CombustionONE
system with a discovery session that includes a site
survey. The initial discovery session will produce:
8
9
1
CombustionONE
TM
Improving and Sustaining the Combustion Asset
For More Information
Contact Us: [email protected]
Yokogawa Corporation of America
2 Dart Road, Newnan, GA 30265-0928
Phone: 770-254-0400 Fax: 770-254-0928
12530 W. Airport Blvd., Sugar Land, TX 77478
Phone: 281-340-3800 Fax: 281-340-3838
http://www.yokogawa.com/us
Yokogawa Canada, Inc.
Bay 4, 11133 40th Street SE, Calgary, AB T2C 2Z4
Phone: 403-258-2681 Fax: 403-258-0182
http://www.yokogawa.com/ca
Yokogawa de Mexico, SA de CV
Av. Urbina No. 18
Fracc. Parque Industrial Naucalpan
Naucalpan de Juarez
Estado de México, C.P. 53489
Phone: (55) 5955-7400 Fax: (55) 5955-7417
http://www.yokogawa.com/mx
Represented by:
TITLE: INNOVATIVE DESIGN OF FCC REGEN FLUE GAS CATALYST
SAMPLER
AUTHORS: KAVYASREE K.V.S.K, VEDULA RAGHU KUMAR, BINOY DAS,
BHASKARJYOTI BARUAH, A.T. NAIDU AND B. BALAGANGADHARAM
COMPANY: HINDUSTAN PETROLEUM CORPORATION LIMITED – VISAKH
REFINERY
ABSTRACT
HPCL – Visakh Refinery is having two FCCUs licensed by M/s Exxon and M/s UOP. The
Regenerator in FCCU-I is equipped with six set of cyclones (six Primary and six Secondary) and
FCCU-II Regenerator is equipped with two set of cyclones (two Primary and two Secondary).
Regular monitoring of cyclones performance is one of the key factors in FCCU performance
monitoring. Periodic data base on Particle Size Distribution (PSD) of fines from FCC
Regenerator cyclones outlet is of high importance.
In Visakh Refinery, Regenerator flue gas catalyst (fines) samplers were installed and
commissioned in both FCCUs at Regenerator outlets. The design was made in-house in a simple
and innovative way. The field modification was carried out online without taking shutdown.
The collection of fines sample from the flue gas catalyst sampler is simple and can be carried out
on periodic basis. Representative sample can be obtained through the sampler, PSD of which
gives correct indication of the performance of the cyclones with time, change in catalyst
properties (if any), etc.
This paper highlights the design, installation and commissioning procedure of the catalyst
sampler implemented at Visakh Refinery.
1. Introduction:
Catalyst losses from FCCU Reactor and Regenerator is a recurring operational issue which
effects Refinery profitability significantly.
Trouble shooting high catalyst losses and
recommending remedial measures is an equally challenging task to FCC Engineers.
One of the highly established and reliable methods for troubleshooting high catalyst losses is by
way of analysis of catalyst Particle Size Distribution (PSD). PSD of fresh catalyst, equilibrium
Page 1 of 5
catalyst and analysis thereafter is a regular practice. Standard charts are available to analyse the
reasons for catalyst loss (attrition, poor cyclone performance, etc.) based on APS and PSD of the
catalyst.
Further detailed analysis includes monitoring of PSD and APS of catalyst lost through the
Reactor and Regenerator cyclones. PSD analysis of the catalyst lost through Reactor side can be
analysed by way of separating out catalyst from Main Fractionator bottoms sample.
Collection of representative sample of catalyst lost from Regenerator side is not standardized in
the design of FCCUs. For FCCUs which are not equipped with tertiary cyclone separator design
or Flue gas scrubber, representative ways of sample collection of catalyst loss through
Regenerator is to be looked in to.
To facilitate the same, Regenerator flue gas catalyst samplers were designed in-house at HPCL
Visakh Refinery and are installed and successfully commissioned in both FCCUs.
2. Details of the Catalyst sampler:
The Engineering details of the catalyst sampler installed in HPCL-VR FCCUs is as given below:
 The flue gas catalyst sampler was placed on the top platform of Regenerator.
 Flue gas inlet to the sampler is from the upstream of DDSV.
 Flue gas outlet from the sampler is let to Atmosphere at safe height.
 Inlet and outlet line sizes are of same size (1”).
 Flue gas flow through the sampler is kept continuous.
 The design temperature and pressure of the sampler are same as that of the Regenerator
operating conditions (Temp: 760 degC and Pressure: 1.6 kg/cm2g).
 Sampler metallurgy is SS. Baffle plates of 18-21 mm thickness were provided to suit for
high temperature operation.
Page 2 of 5
 Brief Process Flow Diagram is given in Fig #1:
Catalyst
sampler
Regen
Orifi
ce
Cha
mbe
r
Flue gas sample
Regenerator
point
 Typical sketch of the sampler is given in Fig #2.
Side view of sampler:
24 inch
Baffle plate
One inch inlet
One inch
Outlet to vent
5 inch
6 inch height
8 inch
8 inch
8 inch
Baffle plate
Top view of sampler:
24” dia
Page 3 of 5
 Catalyst sampler has three compartments separated by baffle plate (two baffle plates).
 Baffle plate is provided with a hole (aperture). Thus, when flue gas flows through the
baffle holes, entrained catalyst particles are settled at the bottom of the sampler and flue
gas passes through it.
 Top lid is provided with proper sealing so that there will not be any leakage of flue gas
from sample collection box. Hinges and bolts arrangement has been provided for the lid
to enable opening of the lid for collection of the sample.
 Photograph of the catalyst sampler is given in Fig # 3.
3. Advantages of the catalyst sampler:
 Easy sample collection
 Zero maintenance
 Sample collected is representative of Cyclones performance over a period of operation
Page 4 of 5
4. Operating data:
 Catalyst collected in the flue gas catalyst sampler is being collected once in a month and
are being analysed for PSD.
 Few results of PSD analysis are provided below:
0-20
wt%
0-40
wt%
0-80
wt%
APS
µm
12.5
42.6
84.9
44.9
Sample
No.
01
0-10
wt%
5.1
02
12.52
46.9
93.6
100.0
20.8
03
8.32
36.3
92.3
100.0
23.2
Basis above analysis, following conclusion was made:
 Sample # 01:
This sample was during high catalyst loss period prior to unit turnaround. Sample
analysis indicates that APS is high indicating that there could be a hole in the
Regenerator cyclone. Same has been confirmed during checking of the cyclones.
 Sample # 02:
This sample was collected during high cat loss problem through the unit. Sample analysis
indicates that 0-10 wt% are higher indicating attrition of the catalyst.
 Sample # 03:
This sample was collected post turnaround and repair of the cyclones. The analysis
indicates normal and high efficient cyclone performance.
5. Conclusion:
Regenerator flue gas catalyst sampler installed at HPCL-VR is a robust, simple, innovative and
successful design that has made catalyst sample collection easier enabling improved
troubleshooting catalyst losses from Regenerator.
Page 5 of 5
TITLE: PROCESSING EXPERIENCE OF HIGH ACID CRUDE –
CORROSION CONTROL STRATEGIES
AUTHORS: A.V.S.KAUSHIK, BHASKARJYOTI BARUAH, A.T. NAIDU AND B.
BALAGANGADHARAM
COMPANY: HINDUSTAN PETROLEUM CORPORATION LIMITED – VISAKH
REFINERY
Contact email: [email protected]
ABSTRACT
Widening of crude basket by processing opportunity crudes plays a significant role in GRM
improvement of a refinery. Processing of high TAN crude as opportunity crude is a challenge for
refiners due to associated corrosion risks.
The crude basket of HPCL – Visakh refinery has more than 100 crudes of different origins,
having wide ranges of physical & chemical properties and varied yield patterns. Processing of
High Acid Crude (HAC) as opportunity crude to maximize profit is a regular practice at the
refinery.
Visakh refinery has three integrated Atmospheric and Vacuum Distillations Units, with a total
crude processing capacity of 8.3 MMTPA. The units were originally not equipped with facilities
to process high TAN crudes. However, as a part of HAC processing plan, systematic study has
been carried out and the modifications required to process high TAN crudes were taken up in a
phased manner in one of the three CDU/VDUs (CDU/VDU-II). Some of the modifications
include change of vacuum column packing metallurgy, installation of corrosion monitoring
facility, HAC chemical dosing facilities, etc.
By way of above modifications (low cost revamp), the unit TAN handling limit has been
increased from less than 0.5 to 1.1 mg KOH/gram on crude oil. Opportunity high TAN crudes
with TAN up to 1.5 are being processed subsequently with ease by limiting the TAN of final
crude blend being processed in the unit to 1.1 and a chemical treatment program.
This paper highlights the details of the modifications carried out in the unit, identification of
strategic locations for corrosion monitoring equipment installation, chemical treatment &
monitoring followed during HAC processing, challenges faced, etc.
1.0
INTRODUCTION
HPCL-Visakh Refinery has three Crude - Vacuum Distillation Units with nameplate capacities
of 1.8, 3.1 and 3.4 MMTPA respectively. The total refining capacity is 8.3 MMTPA.
All the CDU-VDUs are designed for one high sulphur crude (Middle East Crude) and one low
sulphur Crude (Mumbai High). The units were originally not equipped with facilities to process
high TAN crudes.
The strategy adopted for facilitating processing of High Acid Crude was to first carryout a
feasibility study and based on the study findings, carryout metallurgical changes, implement
corrosion monitoring and chemical injection facilities before actual processing.
2.0
STUDY FINDINGS & PLANT MODIFICATIONS CARRIED OUT:
The study objective was to check feasibility of processing high acid crudes in all of the three
CDU-VDUs for three TAN levels (i.e 0.6, 1.1 & 1.7 mg of KOH/g of oil), assess the risks
associated with each circuit, rate the vulnerable circuits based on risks and identify metallurgical
changes required, monitoring requirements & chemical injection locations.
Risk was classified based on TAN of hydrocarbon streams, operating temperatures,
velocity/turbulence, system metallurgy, sulphur content etc. Though the study was carried out for
all the three CDU-VDUs, it was decided to implement the required modifications in CDU-VDU
2. Impact on downstream units like FCCU and DHDS was also evaluated in the study.
Based on the recommendations of the study, the following modifications were implemented in
CDU-VDU 2:
a) Metallurgical changes:
i. Change of packing metallurgy of LVGO, HVGO and Slop cut from SS410 to SS316 (2.5%
Mo min).
ii. Change of HVGO chimney tray and Slop cut chimney tray metallurgy from SS410 to SS316
(2.5% Mo min).
iii. Change of metallurgy of wash section and top demisters from SS410 to SS316 (2.5% Mo
min).
b) Corrosion monitoring facilities:
 Field Signature Method (FSM) logs: Electrodes were installed at the following locations for
close monitoring of corrosion rates.
i. Atmos furnace Radiation outlet first elbow.
ii. Atmos furnace transfer line.
iii.Vacuum furnace radiation outlet first elbow.
 Electric Resistance (ER) probes and Corrosion coupons: ER probes and corrosion coupons
were installed at the following locations to monitor corrosion rate.
i. Crude outlet of final preheat exchanger in preheat circuit
ii. HD CR & Product draw-off line from crude column
iii. LVGO pumps suction
iv. HVGO outlet of HVGO Vs Crude preheat exchanger
v. HVGO pumps suction
vi. Outlet of Atmos overhead condenser
vii. Outlet of Atmos overhead reflux drum
viii. Outlet of vacuum overhead condenser
Additionally, corrosion coupons were installed in suction of Slop cut pumps and Vacuum residue
pumps.
c) Chemical Injection facilities:
To reduce the corrosion rate during HAC crude processing HAC chemical injection points with
diluents were provided at the following locations:
S.No
Injection point
Diluent
1
Suction of Crude booster pumps
Crude
2
Diesel CR and Product common draw off line from Atmos column
Diesel
3.
3
Suction of RCO pumps
Diesel
4
Suction of LVGO pumps
HVGO
5
Suction of HVGO pumps
HVGO
6
Suction of SR pumps
HVGO
ACTUAL PROCESSING EXPERIENCE:
Post implementation of the above mentioned changes, HAC crudes were procured and processed
in CDU-VDU 2 unit. Following are the details of actual processing experience:
3.1 Crudes processed:
i. Bonga crude (TAN 0.66 mg of KOH/g of oil) was processed along with Quaiboe crude.
ii. Usan crude (TAN 1.5 mg of KOH/g of oil) was processed along with Brass Light crude.
3.2 Activities prior to actual processing:
i. Proper functioning of chemical injection facilities and dilution media was ensured.
ii. Proper functioning of corrosion monitoring facilities was ensured.
iii. High sulphur crude was processed in the unit to ensure a sulphide layer on pipelines to reduce
the impact of corrosion due to naphthenic acids.
iv. HAC chemical injection was started 3 – 4 days prior to actual processing of high acid crudes
for passivating the pipelines prone to naphthenic acid corrosion. HAC chemical forms a
protective layer on the metal surfaces thereby mitigating the corrosion due to naphthenic
acids.
v. Baseline data of feed and product samples for TAN, Fe, Ni, V and Ni before processing of
high acid crude was collected for reference purpose.
3.3 Measures taken during processing of high acid crudes:
i. Analysis of feed and product samples for TAN was done every 8 hrs to monitor the corrosion
rate due to naphthenic acids.
ii. Analysis of feed and product samples for Fe, Ni, V and Ni was done every alternate day to
ensure no variation as compared to base case sample results.
iii. Naphthenic acids tend to form soaps when they are neutralized with inorganic cations, such
as sodium in desalter. This results in formation of stable emulsions in desalter. Continuous
monitoring of desalter interface was ensured and desludging frequency was increased to
ensure no rag layer formation due to naphthenic acid precipitation.
3.4 Challenges faced:

Deterioration in vacuum was observed in VDU due to higher vaporization in Vacuum
furnace. The same resulted in Vacuum furnace limitation as required coil outlet temperature
could not be achieved.

Layering issues in crude tanks were observed during the processing due to density difference
of Usan and Brass Lt crudes leading to receipt of Usan crude predominantly initially.

Higher CLO yield was observed in FCCs due to high Nitrogen content in VGO. CLO yield
was higher than normal by ~7.4 wt% and ~8 wt% during Usan & Bonga crudes processing
respectively.

Thruput limitation was observed during processing of high acid crudes due to high VGO
yield. VGO yield was higher by ~2 wt%.

Drop in preheat temperature by ~5 deg C was observed post processing of Bonga and Usan
crudes.
4. RESULTS AND INFERENCES:

The average TAN results of crude and product streams during high acid crude processing are
as follows:
Crude mix
Crude
Diesel
LVGO
HVGO
SR
60 % Usan + 40 % Brass Lt
1.1
1
1.1
1.3
1.9
80% Bonga + 20 % Quaiboe
0.67
0.45
0.66
0.81
1.08
 No abnormal increase in TAN levels of products was observed during high acid crude
processing.

The average metal content analysis of crude and product streams during high acid crude
processing is given below:
S.No
1
Crude mix
60 % Usan + 40
% Brass Lt
2
80% Bonga + 20
% Quaiboe
Metals
Streams metal content (ppm)
Crude
Diesel
LVGO
HVGO
SR
Iron
8.1
0.65
1.85
2.65
60.3
Nickel
9.7
0.44
0.65
0.82
74.4
Vanadium
2.74
0.08
0.11
0.25
19.2
Phosporous
0.04
0.01
0.02
0.05
0.4
Iron
8.2
0.25
2.21
3.5
23.5
Nickel
5.4
0.12
0.16
1.6
44.65
Vanadium
1.1
-
0.84
1.95
14.9
 Metal factor (10 Ni + V) for LVGO and HVGO was observed to be high Vs FCC design of
6.3 ppm.
5. CONCLUSION:
 Strategy of modification of Vacuum column packings, provision of additional corrosion
monitoring instruments and chemical injection facilities is one of the successful revamp
option for processing of high acid crudes in units which are not originally designed for high
TAN.
 High acid crudes were successfully processed in CDU-VDU 2 while maintaining corrosion
rates within limit (<5 mils per year).
CHALLENGES IN PROCUREMENT OF PLATINUM BASED CATALYST
HINDUSTAN PETROLEUM CORPORATION LIMITED– VISAKH REFINERY (VR)
A. SUDHIR, K.VIJAY KIRAN, A.T.NAIDU, B. BALAGANGADHARAM
Contact e-mail: [email protected]
ABSTRACT
HPCL Visakh Refinery (VR) uses Platinum based catalysts in Naphtha Isomerization Unit (NIU)
& Continuous Catalytic Reforming (CCR) unit in MS Block. The catalyst used in these units
undergo deactivation due to continuous operation and needs to be replaced with fresh catalyst at
the completion of the guaranteed life or loss of catalyst activity, whichever happens early.
Platinum based catalyst is expensive and Platinum metal is the major cost component in the
overall catalyst cost. The metal price is governed by international markets and hence is volatile.
The procurement is generally carried out through catalyst suppliers who usually charge
administrative fee for purchase of Platinum metal.
Various challenges are encountered during the procurement process, viz., matching the approval
for advance payment for purchase of Platinum metal with catalyst manufacturing schedule,
Platinum metal accounting & settlement process which includes testing of catalyst samples for
Platinum metal content by refiners and the related procedures & documentation. Due to nonavailability of adequate indigenous testing facilities, the Platinum metal assay report analysis
carried out by catalyst suppliers is being considered for Platinum metal accounting and
settlement. Another significant challenge is with respect to management of Platinum based spent
catalysts. The Platinum metal content in spent catalyst is either sold to prospective bidder or
given to Platinum metal recovery agency to recover the metal. But there are certain issues with
respect to both the options.
This paper addresses various challenges encountered during procurement of Platinum based
catalysts. Present practice for procurement is considered as baseline scenario and various options
were evaluated against the same. Options like, annual lease of Platinum metal with catalyst
supplier, combining catalyst purchase with recovery of platinum metal from spent catalyst and
direct purchase of Platinum & transfer to catalyst supplier were deliberated and each option was
evaluated.
Typical guaranteed life of these catalysts
varies from 3 to 8 years. The deactivated
catalyst is called spent catalyst.
A. INTRODUCTION
HPCL Visakh Refinery uses Platinum (Pt)
based catalyst in NIU & CCR unit in MS
Block. These units were commissioned in
2009 and are in continuous service. The
catalyst used in these units undergo
deactivation due to continuous operation and
has to be replaced with fresh catalyst to meet
the product yields and properties.
The spent catalyst, though deactivated, has
valuable Platinum which can be recovered
(to the extent of 99% on pills and 98% from
fines) and can either be supplied to catalyst
suppliers for fresh catalyst batches or can be
liquidated
through
Platinum
metal
exchanges.
1
The manufacturing of the NIU and CCR
catalysts requires pure Platinum metal. Only
sponge type of Platinum (99.95% pure) is
suitable for catalyst manufacturing. The
required quantity of Platinum metal needs to
handed over to catalyst manufacture at the
time of catalyst procurement.
B. VARIOUS OPTIONS FOR Pt
METAL PROCUREMENT FOR
CATALYST
Fig 1: Platinum price trends for past 10 years
(source: www.kitco.com)
The various options available for Platinum
metal procurement are as follows:
OPTION-1
Procurement of Platinum
along with the catalyst
(from catalyst supplier)
OPTION-2
Annual lease of Platinum
with catalyst supplier
OPTION-3
Recovery of Platinum from
spent catalyst and transfer
of the metal to catalyst
supplier
Fig 2: Platinum price trends for year 2016
(source: www.kitco.com)
CASE A: Platinum metal procurement
through catalyst suppliers:
1. OPTION- 1: Procurement of Platinum
along with the catalyst (from catalyst
supplier):
The catalyst suppliers procure the platinum
from Platinum suppliers and manufactures
the catalyst. The catalyst supplier charges
refinery for the following:
This is the conventional method being
followed till date at refinery. In this option,
an advance is given to catalyst for
procurement of required quantity of
Platinum metal upfront after placement of
Purchase order. Bank guarantee of
equivalent value is submitted by the catalyst
supplier to refinery.
a. Price of
Platinum metal based on
London PM fix on that day – Price in
USD (US Dollars) /TOZ (Troy Ounce) of
Pt (advance amount is paid to the
supplier and Platinum settlement is
carried out at a later stage as per the
Catalyst Supply Agreement (CSA))
The Platinum metal price is driven by the
international markets and can vary to a large
extent and therefore affects the overall cost
of the procured catalyst. The Platinum
pricing trend for the past 10 years and 1 year
is given below:
b. Commission of trader – 3 USD/ TOZ of
Pt (typical).
c. Administrative charges – 6000 USD
(typical)
2
The total expenditure for refinery is the sum
of a, b & c.
CASE B: Platinum metal procurement
directly from Platinum suppliers
Some of the platinum supplier available are
M/s Johnson Matthey, M/s Haeraeus etc.
Platinum can be procured by refinery from
the Platinum suppliers directly. The
following are the requirements:
On spot pricing (prevailing price on the
Limited order (when no urgency in
If transfer is from a different location,
the charges applicable are - 1
USD/TOZ
2. OPTION-2: Annual lease of Platinum
with catalyst supplier
 In this option, the required quantity of
Platinum metal sponge is leased by the
catalyst supplier to refinery, with an
applicable annual lease charge.
specified date)

Charges for transferring Platinum
from refinery account to catalyst
suppliers (Incase the transfer affects in
The sum of the above three components
under ‘f’ is the total expenditure for
Refinery.
c. Three options for booking/ fixing of price
(pricing) of the catalyst is done:
Average price of Platinum during the
month


b. The required quantity Platinum is to be
booked with Platinum supplier.

Price of Platinum metal based on
London PM fix on that day – Price in
USD/TOZ of Pt
the same country, the charges are Nil)
a. A Platinum account to be opened by
refinery with Platinum supplier (no
charges are levied for opening the account).


 Typical lease periods are one year
from the date of platinum handing
over to catalyst suppliers. Lease is
revised on annual basis.
procurement exists, a particular price
may be intimated to the supplier to buy
when the Platinum cost reaches to a
certain value)
 Annual lease rate can vary from 2% ~
4% of the total platinum cost
depending on the catalyst supplier.
d. Once the Platinum booking is accepted
by supplier, Pro-forma Invoice is sent to
refinery and transfer of funds is done.
Platinum procured is deposited in
refinery account with the Platinum
suppliers.
 Refinery will be required to provide
Bank guarantee on 100% of the
Platinum cost during the lease period.
 The lease rate and other terms and
conditions are agreed in Platinum
Lease Agreement (PLA).
e. Based on the requirement, catalyst is
transferred from refinery account to
Catalyst suppliers account.
 The lease can be fore-closed by
settling the equivalent quantity of
Platinum metal with catalyst supplier.
f. The total expenditure involved in this
method is given below:
 The option of lease is viable if
sufficient spent catalyst (containing
Platinum) is available for Platinum
3
from the time of placement of PO.
Therefore, internal approval may be
required for extending the PO value to
cover for the increased Platinum metal
prices. This further delays the
procurement activity and delivery of the
catalyst.
recovery. However, the following
requirements are to be ensured:
 Refinery has to identify the Platinum
recovery agency and place PO for
metal recovery.
 A detailed agreement with the
Platinum recovery agency is made on
the percentage recovery of the
Platinum, purity of the platinum,
supply of Platinum sponge in “Good
delivery bottles” to the vaults of the
Platinum exchanging companies of the
catalyst supplier.

Due to non-availability of adequate
indigenous testing facilities, the
Platinum metal content analysis carried
out by catalyst suppliers is being
considered
for
Platinum
metal
accounting and settlement.
 Typically, the recovered Platinum
metal is squared off with leased metal
quantity and lease can be fore-closed.

3. OPTION-3: Recovery of Platinum
from spent catalyst and transfer of the
metal to catalyst supplier.
Management of Platinum based spent
catalysts.
The Platinum metal content in spent
catalyst can either be sold to prospective
bidder or given to Platinum metal
recovery agency to recover the metal.
In this option, the Platinum metal sponge
is recovered from spent catalyst either
through catalyst supplier or a Platinum
recovery agency and required quantity of
metal is deposited in catalyst
manufacturer Platinum metal accounts.
The same is used for manufacture of
fresh catalyst on need basis.
However, it is observed that only few
options are available in India for
Platinum metal recovery. Also the
percentage of metal recovery offered by
agencies in India is less compared to the
percentage metal recovery offered by
overseas agencies. Transportation of the
spent catalysts to overseas agencies
requires clearances from Ministry of
Environment and Forests (MoEF) as the
spent catalyst is classified under
Hazardous waste.
C. CHALLENGES FACED DURING
PLATINUM PROCUREMENT
Various challenges that are encountered
during the procurement process are
listed below:

Platinum metal accounting & settlement
process which includes testing of
catalyst samples for Platinum metal
content by refiners and the related
procedures & documentation.
D. METHODOLOGY ADOPTED IN
REFINERY FOR PROCUREMENT
Matching the approval for advance
payment for purchase of Platinum metal
with catalyst manufacturing schedule,
The conventional method, as indicated in
Option-1, is presently being followed at
Refinery in view of the challenges indicated
At times, due to volatility of Platinum
metal prices, the metal prices may vary
4
in Section C. The same is also observed to
be followed by many other Indian
Refineries.
advance
for
procurement.
viii. Catalyst manufacture by Catalyst
Supplier and supply as per the delivery
terms and conditions of PO.
Salient aspects of the Platinum metal based
catalyst procurement are given below:
Preparation of Purchase requisition
indicating the details i.e., catalyst
quantity, Platinum metal quantity etc
ii.
Request for quotation (RFQ) for supply
of Platinum metal catalyst
iii.
Technical
evaluation
&
commercial
iv.
Placement of Purchase order (PO) to
Catalyst Supplier
v.
Finalization of
agreement (CSA)
vi.
Submission of bank guarantee by the
Catalyst Supplier for release of
Catalyst
metal
vii. Release of advance to catalyst supplier
for Platinum metal procurement.
To ensure timely procurement of Platinum
based catalysts, close follow up with inter
departments like Materials, Finance and
catalyst suppliers are required.
i.
Platinum
bids
ix.
Analysis of catalyst samples for actual
Platinum
content
followed
by
exchange of analysis between catalyst
supplier & refinery to ascertain actual
quantity of catalyst used in catalyst
manufacture and estimate balance
quantity of Platinum.
x.
Basis exchange of analysis, settlement
of balance quantity of Platinum metal.
E. BENEFIT
ANALYSIS
VARIOUS OPTIONS:
FOR
A preliminary benefit analysis for
procurement of platinum catalyst with
identified options indicated under section B,
are provided in Table-2.
supply
Table - 2: Benefit analysis for various options
DESCRIPTION
XX-R-XX Catalyst requirement
Platinum requirement for catalyst
manufacture (Pt @0.2 wt%)
Average price of Platinum (assumed)
 Cost of Platinum for procurement
LEASE COST COMPONENTS
 Cost of lease for Platinum @ 4% of
Platinum cost per annum
 Bank guarantee equivalent to cost of
Platinum and the related charges (@10%
per year)
UOM
Kg
Kg
TOZ
USD/TOZ
USD
OPTION-1
Cat procurement
+
Pt procurement
25,660
59.0
1,897.5
OPTION-2
Catalyst
procurement
+Pt Lease
25,660
59.0
1,897.5
OPTION-3
Catalyst
procurement +
Pt recovery
25,660
59.0
1,897.5
1,000
18,97,471.35
-
-
$
%
-
$
75,898.0
-
USD
-
$
1,89,747.0
-
5
DESCRIPTION
 Cost for Platinum metal recovery from
spent
catalyst
(1%
handling
charges+3USD per Kg of catalyst)
Conversion of Platinum metal to
ChloroPlatinic acid (CPA) (required as a
part of catalyst preparation process) (193
USD per Kg)
Cost of the catalyst (excluding Platinum
metal) (assumption)
UOM
OPTION-1
Cat procurement
+
Pt procurement
OPTION-2
Catalyst
procurement
+Pt Lease
OPTION-3
Catalyst
procurement +
Pt recovery
USD
-
$
95,954.0
$
95,954.0
USD
$
10,204.0
$
10,204.0
$
10,204.0
USD
$
8,47,626.0
$
8,47,626.0
$
8,47,626.0
EXPENDITURE FOR PROCUREMENT OF CATALYST
First year
Second year
Total expenditure for procurement of
catalyst at the end of two years
(1
USD=68 ₹)
USD
₹
USD
₹
$ 27,55,302.0
₹ 18,73,60,558.0
Not Applicable
USD
₹
$
27,55,302.0
₹ 18,73,60,559.0
₹ Crores
18.74
$ 12,19,431.0
₹ 8,29,21,355.0
$ 2,65,646.0
₹ 1,80,63,927.0
$
14,85,077.0
₹ 10,09,85,282.0
8.3 (1-year lease)
10.1
(2
years
lease)
$
9,53,785.0
₹ 6,48,57,427.0
Not Applicable
$
9,53,785.0
₹ 6,48,57,427.0
6.49
OPTION 1: Procurement of catalyst along with Platinum metal.
OPTION 2: Procurement of catalyst with Platinum metal quantity on lease.
OPTION 3: Recovery of Platinum metal from spent catalyst and transfer to catalyst supplier.
Based on the benefit analysis for each
option, following are the observations:
4. If the Platinum recovery from spent
catalyst is expected to take some time,
catalyst procurement with lease quantity
of Platinum metal may be a feasible
option (Option 2). The lease period may
be considered for 1 year. During the
lease period, the Platinum recovery
process can be initiated such that
recovered Platinum can be settled with
catalyst supplier by the end of lease
period.
1. The Option-1, which is in practice
presently, is involving a very high
expenditure against the other two
options (18.7 crores vis a vis 10 (Option
2; 2 year lease), 8.3 (Option 2; 1 year
lease period) and 6.5 crores for Option
3).
2. The Option-3 i.e. Platinum recovery and
procurement of catalyst is the best cost
economic option. (The catalyst cost
(excluding Platinum is 5.76 crores))
F. CONCLUSION:
3. It is observed that instead of
procurement of Platinum, it is beneficial
to recover the Platinum metal from
available spent catalyst as Platinum and
utilize the same for future catalyst
procurement.
In view of the above, it is observed that with
proper planning and platinum metal
management, especially at the time of
catalyst procurement, the overall cost of the
platinum catalyst can be reduced by one
6
third. It is therefore imperative to evaluate
various options to reduce the cost of
platinum catalyst at the time of procurement
and implement the most cost effective
option subject to the suitability.
7
GRM IMPROVEMENT - MAXIMIZING DIESEL & HYDROGEN IN
REFINERY WITH MINIMIUM CAPEX
AUTHORS: B. GOPI, G V MADAHV, K. VIJAY KIRAN, A.T. NAIDU, B. BALAGANGADHARAM
COMPANY: HINDUSTAN PETROLEUM CORPORATION LIMITED – VISAKH REFINERY
Contact e-mail: [email protected]
ABSTRACT
The demand for diesel is increasing day by day globally. To meet the increasing diesel demand,
refineries in the world are exploring various opportunities to maximize distillate yields. To maximize
diesel yield with existing refinery configuration and without incurring additional capital cost is a
challenge. In HPCL-Visakh Refinery (VR), Heavy Naphtha (HN) stream from Crude Distillation Units
(CDU) is predominantly blended with diesel during E-III regulation. However, with Euro-IV diesel
requirement in place, this medium/high sulfur HN is rejected and is eventually exported resulting in
loss to refinery. The same is the scenario for many refineries globally. Hydro treating high/medium
sulfur HN stream was identified as a prospective method for minimizing exports. However,
establishing a new unit involves huge investment & cost. Therefore, feasibility of utilization of idle
assets for processing HN was carried out. Also, to improve hydrogen yield and minimize operating cost
of Hydrogen Generation Units (HGU), feasibility of using Light Naphtha (LN) from Naphtha Hydro
treating Unit (NHT) in place of sour Straight Run Naphtha (SRN) was explored. This innovative twin
pronged strategy resulted in hydrogen yield increment in HGUs and reduction in operating cost by
stopping Pre-Desulfurization Section (PDS) of HGU, reduction in naphtha/steam consumption. The
idle assets created i.e. PDS, was utilized to hydro treat medium/high sulfur HN. The implementation of
these schemes required additional piping requirement and modification of operation conditions both in
HGU/PDS units. Implementation time was 4 months (for HGU) and 1 month for PDS unit. The hydro
treated PDS HN is routed to diesel and the benefits of 3.8 million USD per year is achieved.
Additionally, hydrogen yield improvement of 2% is achieved in HGU.
A. INTRODUCTION
HPCL-VR is presently configured with two
Hydrogen Generation Units (HGUs) namely
DHDS-HGU and DHT-HGU to meet the
hydrogen demand of the refinery. The units
were installed as a part of DHDS and DHT
projects in 2000 and 2014 respectively. M/s
Technip is the process licensor of both HGUs.
The design hydrogen generation capacities of
DHDS-HGU and DHT-HGU units are 18320
TPA and 36000 TPA respectively.
The Hydrogen generation plants are based on
Steam-Naphtha Reforming process and are
governed by the following reactions:
Reforming:
CnHm + nH2O → nCO + (2n + m)/2 H2
CH4 + H2O
→ CO + 3H2
C + H2O
→ CO + H2
Shift Conversion:
CO + H2O
Page 1 of 8
→ CO2 + H2
The HGU consists of Pre-Desulphurization
Section (PDS), Final Desulphurization Section
(FDS) and Reformer section. Further
conversion of CO to H2 is achieved with shift
reaction section. The final purification of the
process gas is carried out by Pressure Swing
Adsorption (PSA). DHT-HGU is configured
with the advanced technology viz. ultra
desulphurization in FDS, pre-reformer, low
temperature shift conversion etc. to improve
hydrogen recovery.
The SRN from CDUs is taken as feed to PDS
section, where sulphur in feed is reduced to 10
ppm and it is further desulphurized to less than
0.1 ppm in FDS section to meet the
requirements of reforming section. SteamNaphtha reforming reactions takes place in
reformer, further hydrogen conversion is
achieved in shift reaction section and finally
sent to PSA for achieving 99.99% pure
hydrogen.
To increase hydrogen yield and maximize
diesel production with minimum CAPEX, the
following innovative twin pronged strategy is
developed:


Using low sulfur LN (from NHT) in
place of SRN as feed to HGU reformer.
This resulted in discontinuing the
operation of PDS section and also
increased hydrogen yield by 2%.
Processing of HN (from CDUs) in PDS
section of HGU, which is routed to E-IV
diesel thereby maximizing diesel yield.
This paper provides detailed methodology
adopted for evaluation and implementation of
the above mentioned strategy and the benefit
achieved.
B. STRATEGY
FOR
INCREASING
HYDROGEN AND MAXIMIZATION
OF DIESEL YIELD.
1. CHANGEOVER OF REFORMER FEED
FROM SRN TO LN IN DHDS-HGU:
DHDS-HGU was commissioned in the year
2000. SRN was the only available feed for
HGU and the same is considered in the design.
Naphtha Hydro treating Unit (NHT) was
commissioned subsequently in 2009 as a part of
Visakh Refinery Clean Fuels Project (VRCFP)
to meet Euro-III MS specifications. Post
commissioning of NHT, SRN from CDUs is
routed to NHT from which LN and HN with
ultralow sulfur (less than 0.5 ppm) are
produced. LN is routed to NIU partly and
balance to MS pool or Naphtha pool depending
upon MS blend specifications. Entire HN from
NHT is routed to CCR for producing reformate.
NHT yields about 45 to 50% i.e. around 115120 m3/hr of LN whereas NIU is designed only
for 42 m3/hr of LN and therefore the balance
LN is routed to MS pool or Naphtha. Of all the
streams routed to MS pool, LN is the least
chosen stream in view of low RON and is
usually rejected to Naphtha. Therefore, various
options to upgrade LN were explored.
The average properties of LN were analyzed to
evaluate the suitability of processing as HGU
Reformer feed. The major parameters analyzed
and are tabulated below:
TABLE-1: PARAMETERS OF SRN VS LN
PARAMETER
Source
Sulfur, ppm
End point (EP), 0C
Paraffins, %Vol
Density, kg/m3
SRN
CDUs
500 to 1000
160
70
697
LN
NHT
0.1
95
80
670
It is evident from the above table that LN has
low sulfur, low end point and high paraffins as
Page 2 of 8

compared to SRN stream. The following are the
envisaged benefits due to the above
characteristics of LN:

Ultra low sulfur of LN doesn’t
necessitate operation of PDS section.
(reduction in opex of HGU)

Higher paraffins in LN results in higher
hydrogen yield. (estimated to be ~3%
increase in hydrogen yield)
Lighter feed (lower FBP of LN) results
in higher catalyst life in view of reduced
coke formation.
The paraffin content and final boiling point
of LN and SRN are shown in Figures 1 & 2
respectively.
FIG.1 SRN vs LN PARRAFIN CONTENT
90
PARRAFIN CONTENT (VOL%)
85
80
75
70
65
60
LN
SRN
FIG.2 SRN vs LN FBP
180
160
FBP (DEG C)
140
120
100
80
60
40
20
0
SRN
LN
Page 3 of 8
Based on the above analysis, it is proposed to
process LN as feed to HGU in place of SRN.
Therefore, a detailed in-depth analysis of LN
stream properties, feed contaminants, impact of
LN on reformer catalyst, review of reformer
operating conditions, adequacy of downstream
equipment viz., shift reactor and PSA were
carried out. Also the logistics related to
sourcing of LN to HGU was also reviewed.
A scheme was developed and following are the
salient features of the scheme:




Provision of piping from LN (from
NHT) to FDS feed tank.

The major operating conditions for SRN and
LN cases at turndown and 65% load is
tabulated below for comparison.
TABLE 2: OPERATING CONDITIONS AT
TURNDOWN CAPACITY
PARAMETER
Reformer outlet temp. 0C
802
799
HT shift inlet temp. 0C
312
311
HT shift outlet temp. 0C
380
372
PSA Inlet CH4 (vol %)
2.01
2.0
PSA Inlet CO (vol%)
2.01
2.23
SRN
LN
4960
4965
689
660
Reformer inlet temp. C
508
491
Reformer outlet temp. 0C
817
812
HT shift inlet temp. 0C
314
321
HT shift outlet temp. 0C
396
397
PSA Inlet CH4 (vol %)
3.5
2.4
PSA Inlet CO (vol%)
3.3
2.8
Feed Rate (Kg/Hr.)
Density (Kg/m3)
0
SRN is continued as fuel to reformer.
OPERATING CONDITIONS:
LN
PARAMETER
PDS of HGU is shut down and kept
under inert condition. The operation of
PDS
is envisaged
during the
turnaround/downtime period of NHT.
PERFORMANCE
ANALYSIS
POST
SWITCH OVER FROM SRN TO LN
SRN
TABLE 3: OPERATING CONDITIONS AT 65%
CAPACITY
Existing FDS feed pumps and piping is
found to be adequate for processing LN.
The proposed scheme was implemented within
4 months complying to the Management of
Change procedure. The capex involved in
implementing the scheme was ₹10 lakhs.
PARAMETER
It is found from the above tables that no
appreciable change in operating conditions is
made during LN processing.

YIELD AND NAPHTHA/H2 RATIO:
The hydrogen yield and Naphtha to hydrogen
ratio was calculated for both the cases and the
same is tabulated under Table 4 & 5.
TABLE 4: HYDROGEN YIELD WITH SRN AND
LN FEEDS
HYDROGEN YIELD (WT%)
SRN
LN
Feed Rate (Kg/Hr.)
2740
2745
Density (Kg/m3)
682
655
Reformer inlet temp. 0C
492
491
FEED TYPE
DESIGN
ACTUAL
SRN
25.5
25.3
LN
----
27.2
Page 4 of 8
TABLE 5: NAPHTHA TO H2 RATIO
INFERENCES:
NAPHTHA TO H2 RATIO (kg/kg)
FEED TYPE
DESIGN
ACTUAL
SRN
3.92
3.92
LN
----
3.65
The trend of hydrogen yield with SRN and LN
is shown in Figure. 3

An increase in hydrogen yield of 2 wt % is
achieved with the feed switch over from
SRN to LN.

The above improvement was achieved with
minimum capex and without major changes
in operating conditions.

Considering the above benefits, the scheme
was extended to the newly commissioned
DHT-HGU
FIG.3 SRN vs LN H2 Yields (wt%)
28
27.3
HYDROGEN YIELD (WT%)
28
27.5
27.3
27.1
27.1
26.9
27.5
26.8
26.5
26.5
27
26
26
25.3
24.8
25.1
26
25.4
25.7
25.3
25
25
24
23
1
2
3
4
5
6
SRN
2. HN PROCESSING IN PDS:
Heavy Naphtha (HN) stream from CDUs was
predominantly blended with hydrotreated diesel
during E-III regulation. However, with Euro-IV
diesel regulation in place, this medium/high
sulfur HN stream was rejected and was
eventually exported resulting in loss to refinery.
Also, processing high/medium sulfur HN in
existing DHDS/DHT along with sour diesel
was not feasible due to limitations in stabilizer/
fractionation sections.
7
8
9
10
HN
View above, hydro treating of high/medium
sulfur HN stream was identified as a
prospective method for minimizing exports.
However, installing a new unit involves huge
capex. Feasibility of utilizing idle assets for
processing HN was checked. The PDS section
of HGU units were idle since the feed
changeover of HGU from SRN to LN.
Therefore, the PDS section of HGUs was
reviewed for processing HN.
Page 5 of 8

FEED DESIGN BASIS COMPARISON:
Critical parameters of HN blend generated in
CDUs are compared with the PDS design feed
specifications.

Provision of piping for off-spec Naphtha
from PDS battery limit to HN feed tank.
The above scheme is implemented within one
month complying to standard refinery
management of change procedure.
HN Blend sulfur is within the limits (800
ppm vs 1375 design max.
Blend EP is marginally higher. (186 vs 176
design max)
Blend aromatics is within the limit (14
vol.% vs 15 vol.% design max)
Blend Olefins is within the limits (0.2 vol%
vs 18vol% design max)
Total metal content was within the design.
CHALLENGES DURING START-UP &
STABILIZATION:
All the parameters are within the design except
the distillation (EP) which is slightly higher
than the design, necessitating modifications in
the operating conditions of the stripper for H2S
removal.
Further to the feed design basis comparison,
PDS unit is reviewed for the following for HN
processing:
The root cause for the above challenges is
identified as the high EP of HN (215 deg C
against design of 176 deg C).








Suitability of the catalyst
Rotating equipment adequacy
Pre-heat circuit and heater adequacy
Operating conditions and stream routing
Based on the analysis, PDS is found to be
suitable for processing HN. Hence, a scheme is
proposed to route high/medium sulfur HN from
CDUs. The following are the major
modifications of the scheme:



Conversion of redundant JBO tank to HN
feed tank.
Provision of piping from HN feed tank to
suction of Naphtha pumps.
Provision of piping from PDS rundown to
DHT diesel rundown line.
The following challenges were encountered
during startup and stabilization period:
1. High PDS product naphtha sulphur and
presence of H2S.
2. Frequent plugging of feed filters
3. Heater limitation.
TROUBLESHOOTING & ACTIONS
TAKEN:
 Simulation study of stripper system with the
actual distillation of HN was carried out to
estimate the required operating conditions
for high EP operation. The estimated
conditions are slightly higher than the
design conditions.
 A detailed study on mechanical integrity of
the stripper system and its related loops was
carried out for operation at higher than
design temperature.
 Stress analysis of stripper piping system
including supports was carried out by an
external agency.
 Particulate analysis of feed and product is
carried out.
Based on the above studies, the following
recommendations are implemented:
Page 6 of 8
 Minor modifications in the support gaps of
piping related to stripper feed and
feed/bottom exchanger are carried out.
 Stripper operating conditions are revised
(stripper bottom temperature increased
from 180 to 215 deg C and operating
pressure decreased from 6.5 to 3.5 kg/cm2
g).
 Reactor inlet temperature was optimized
(reduced from 300 to 285 deg C).
 Heater operation is optimized to overcome
limitation (excess air, burner cleaning etc.,)
 Frequency of feed filter elements
replacement increased.
 The stripper off gas estimated at lower
pressure operation is ~ 10 kg/hr and the
purge gas being flared from recycle gas
loop is estimated to be ~11 kg/hr is routed
to flare against normal routing to fuel gas
system due to low pressure operation.
The above recommendations are implemented
and the unit is operated with the revised
conditions
INFERENCES:


The unit is operated with HN at design
capacity on a consistent basis and the
entire product is blended with Euro- IV
diesel.
The average feed sulfur of HN
processed in PDS is 760 ppm and
achieved a product sulfur of 5 ppm. The
distribution of feed and product sulfur
shown
in
the
Figure
4.
1200
50
45
40
35
800
30
600
25
20
400
15
PRODUCT SULPHUR
FEED SULPHUR
1000
10
200
5
0
0
Feed Sulphur
C. CONCLUSIONS:

The twin pronged strategy of
changeover of SRN to LN as feed to
HGU coupled with the usage of idle
asset
(PDS)
to
process
high
sulfur/medium sulfur HN from CDUs
has resulted in an estimated benefit of
₹25 crores per annum.
Product Sulphur

In addition to the benefits of blending
the hydrotreated HN in E-IV diesel, an
improvement of hydrogen yield of ~2%
is achieved.

A capex of ₹20 lakhs is utilized for the
implementation of the above schemes.
Page 7 of 8

These schemes were implemented
quickly and benefits were realized.
Page 8 of 8
satimataji
RTM: 2017
State-of-the-art methodology for determination of Chloride in
hydrocarbon streams using modified extraction and potentiometric
determination.
Dr Y S Jhala*, Dr Ashutosh Mishra, Dr H K Singh, Dr B R Panda, A K Nath, D Diraviyum,
S Sarkar, D Chakraborty,
Indian Oil Corporation limited, Gujarat Refinery, Vadodara
[email protected]
Abstract
Organic chlorides can cause corrosion in pipelines, valves and condensers, and catalyst
poisoning. The hydrocarbon processing industry and others have been affected by the
potential damage created by these substances. Organic chloride species are potentially
damaging agent for refinery processes. Hydrochloric acid can be produced in hydro treating
or reforming reactors and the acid accumulates in condensing regions of the refinery.
Unexpected concentrations of organic chlorides cannot be effectively neutralized which turns
to damage.
Organic chlorides in crude oil during distillation concentrate in heavy naphtha fractions and
can cause extreme damage if not properly analyzed and reported. Timely and accurate
reporting of organic chlorides before refining helps refiners to take proper corrective action.
Even very low chlorine concentrations in crude oil are detrimental, no matter whether they
come from inorganic salts or organic compounds. Obviously a fast and reliable determination
of the chlorine concentration in crude oil and related matrices is beneficial for any refinery.
The available methods are well suited with lighter cuts and they also require costly
equipment.
Present work deals with development and validation of new method for determination of
chloride content from Naphtha cut to VGO cut of refinery. Precision and accuracy of new
methodology has been stabilized by statistical calculation.
*Corresponding Author
Satimata
Introduction
Chloride content in hydrocarbons is an important characteristic of oil quality. Organic
chloride species are potentially damaging agent for refinery processes. Hydrochloric acid
can be produced in hydrotreating or reforming reactors and the acid accumulates in
condensing regions of the refinery. Unexpected concentrations of organic chlorides cannot
be effectively neutralized and results can be leakage and also leads to catastrophic failure.
Naphtha is often used as a solvent for petrochemical feedstocks, as well as for the
production of natural gas. Chlorides are frequently occurring contaminants in naphtha and
often exceeding at desired concentration. While some chlorides may be naturally found in
crude oil, they can also be introduced during cleaning processes or use of chemicals during
desalting process.
Chlorides can be present as organic chlorides, organic compounds with a carbon–chloride
bond, or undesaltable chlorides, which are organic compounds that cannot be removed
during a salt separation process. High chloride levels are often reported in refining
equipment, and are undesirable as they can cause scaling and ultimately damage to the
equipment.
However, during the hydrotreatment process, chlorides that are present in naphtha and
crude oil can lead to the formation of hydrochloric acid, which can then corrode equipment.
To minimize corrosion and scaling, chloride levels should be kept below 1 ppm. Chlorides
can be detected and monitored in naphtha and crude oil through a reduction with sodium
biphenyl and subsequent potentiometric titration. This chloride is not permanently bound to
the support and is therefore sent out with all product streams. Effects of chloride in
downstream operations can be:

Corrosion of austenitic stainless steels

Ammonium chloride salting - corrosion and pressure drop issues

Corrosion of downstream piping/equipment
More recently, a growing concern for many refinery operators has become the removal of
organic chloride species. These compounds are less easy to detect and measure and also
are less readily ad- or ab-sorbed. The effectiveness of the available chloride detection
techniques limited wrt to type of sample and phase of sample.
Satimata
In view of above accurate and trace level monitoring of Chloride is highly required. There are
ASTM and UOP methods available for Organic Chloride Determination in Crude Oil and
Naphtha Fractions. The available methods Method having some limitations
like use of
Sodium Biphenyl Reduction which is having low stability period and second one is use of
Combustion and Microcoulometry which requires costly equipment. Combustion and
Microcoulometry method is only for lighter ends like Naphtha.
The aim of the present work is to develop a new state of art Methodology for the detection of
Organic and Total Chloride at law level in Naphtha to VGO Streams by using Sodium Metal.
In this work we standardize a process by using Sodium based extraction and Potentiometric
determination of Chloride in different hydrocarbon Streams. The chemistry behind this is
very simple i.e. When a sodium atom transfers an electron to a chlorine atom, forming a
sodium cation (Na+) and a chloride anion (Cl-), both ions have complete valence shells, and
are energetically more stable. The ionic bonded NaCL is soluble in water and the chloride
containing aqueous phase is acidified, concentrated, acetone is added and the solution
titrated potentiometrically
Experimental
A. Sample Preparation:
The method is useful for Naphtha and Higher Hydrocarbons. High Viscous hydrocarbons like
VGO can be diluted in Benzene or Toluene. The Sample size depends on Chloride
Concentration.
B. Procedure of Testing:
Take required sample in separating funnel. In case of H2S traces, Wash the Sample with
KOH to remove H2S and then wash with distilled water. Take 30 to 50 gm. sample in round
bottom flask add 0.80 gm sodium metal in toluene ( dry sodium metal).Initially heating starts
at very slow until sodium metal dissolved. Remove the Flask and cool it to Room temp.
Transfers the solution into 250 ml beaker and add 30 % Nitric acid, concentrate the
content up to 75% by boiling. To separate the aqueous layer, again acidify it with dilute (1:7)
Nitric acid and shake it well. Drain out lower aqueous layer in 50 ml volumetric flack. Repeat
the step twice to get complete extraction. Same way carried out blank preparation also. Take
10ml aqueous layer and add 70 ml acetone and titrate the solution potentiometrically with
standardized 0.1- or 0.01-M silver nitrate using glass versus silver rod electrodes. The blank
is titrated in similar way.
Satimata
Validation of Method
The Method Validated by comparison of same by standard Method ASTM D 4929 and
results is under repeatability. To further validation of extraction procedure and evaluation
procedure a Standard Sample of Chlorobenzene has been used and recovery of standard
also established the method.
Following few data clearly establish the Method:
Sample
D4929
New Method
Naphtha
2.3 mg/kg
2.1 mg/kg
Naphtha
0.88 mg/kg
0.90 mg/kg
Standard Chlorobenzene 10.0 ppm
9.9 mg/kg
10.2 mg/kg
Standard Chlorobenzene 1.0 ppm
1.0 mg/kg
1.03 mg/kg
Inorganic Chloride can also done by similar extraction without addition of Sodium Metal
Conclusion
The developed methodology was able to detect Chloride in various hydrocarbon samples
with good repeatability. This system enables us to identify and control gas leakage from
units which further improve cooling water performance.
Acknowledgements
Authors are thankful to management for permission to publish this paper. A hearty thanks to
QC colleagues of Vadodara and Technical Services for their advice and assistance provided
during the study in the laboratory.
References
1.
ASTM D 4929.
2.
Primer on Organic Chlorides and Their Control , S.A. Lordo . Nalco Energy Services.
3.
Na e international Pu li ations” Effect of non-extractable chlorides on refinery
corrosion and fouling
Satimata
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