School of
Chemical
Engineering
SKKC 4143 PLANT DESIGN
2018/2019-SEM 1
FINAL REPORT
PROPYLENE PRODUCTION PLANT
LECTURER
DR ZARINA MUIS
DESIGN TEAM (GROUP 6)
NO.
TEAM MEMBERS
1. DINESH A/L SIVARAJU
2. SHAMIMY NAADIAH BINTI MOHD SHOKRI
3. NUR ZAHIRAH BINTI HASLI
MUHAMMAD AIMAN AMIRUDDIN BIN MHD
4.
KAMAL
MATRIC NO
A15KK0027
A11KK0140
A15KK0111
A15KK0220
ii
ACKNOWLEDGEMENT
First and foremost, we would like to express our gratitude to our beloved
lecturer of Plant Design subject, Dr. Zarina binti Ab Muis for her guidance and advice
in supporting us to complete this project. This report would not been completed
without the help from her.
Finally, we hope this report will be helpful for readers in studying about the
process of propylene production.
iii
TABLE OF CONTENTS
TITLE
ACKNOWLEDGEMENT
PAGE
ii
TABLE OF CONTENTS
iii
LIST OF TABLES
vi
LIST OF FIGURES
vii
CHAPTER 1: INTRODUCTION
1
1.1
Background of Propylene
1
1.2
Production of Uses of Propylene
2
1.3
Propylene Manufacturing
3
1.4
Process Flow Diagram
5
1.5
Market Survey
1.6
6
1.5.1 Introduction
6
1.5.2 Production of Propylene
7
1.5.3 Propylene Consumption
8
1.5.4 Local Outlook for Production of Propylene
10
1.5.5 Market Prices of Propylene
11
Process Screening
12
1.6.1 Gross Profit
12
1.6.2 Factors Affecting Screening Process
14
1.6.2.1 Temperature
1.6.2.2 Pressure
15
1.6.2.3 Safety
15
1.6.2.4 Environment
1.7
Site Location
1.7.1 Water Tariff
1.8
14
15
16
17
1.7.2 Electricity Tariff
19
Conclusion
19
iv
CHAPTER 2: PROCESS CREATION AND SYNTHESIS
20
2.1
Introduction
20
2.2
Sources of Raw Material and Its Specification
20
2.3
Physical and Chemical Properties
2.4
Synthesis Steps
2.5
Manual Mass Balance Calculation
2.6
21
22
37
2.5.1
Overall Mass Balance
37
2.5.2
Mass Balance for Mixer
38
2.5.3
Mass Balance for Reactor
39
2.5.4
Mass Balance for Separation Unit
40
2.5.5
Mass Balance for Separation Unit 2
41
2.5.6
Mass Balance for Separation Unit 3
42
2.5.7
Summary Table Mass Balance
43
Manual Energy Balance Calculation
44
CHAPTER 3: PROCESS SIMULATION AND HEAT INTEGRATION
3.1
3.2
47
Percentage Difference between Manual Calculation and
HYSYS Calculation
47
Heat Integration
49
3.2.1
Calculation of FCp
49
3.2.2
Process Energy Integration
55
3.2.3 Algorithm Table
55
3.2.4 Heat Exchanger Network
57
3.2.5 Comparison of Performance Before and After Heat Integration
59
3.2.6 Process Flow Diagram Heat Exchanger Network
60
CHAPTER 4: PROCESS OPTIMIZATION
4.1
Process Optimization
61
CHAPTER 5: EQUIPMENT SIZING AND COSTING
5.1
65
Sizing and Costing of Equipment
5.1.1
Oleflex Reactor
65
65
v
5.1.2
Pump
5.1.3
Cryogenic Separator
5.1.4
Distillation Column
5.1.5
Heat Exchanger
5.1.6
Cooler
5.1.7
Heater
66
67
69
75
83
85
CHAPTER 6: TOTAL CAPITAL INVESTMENT
6.1
Total Capital Investment
6.1.1
88
88
Estimation of Total Capital Cost Investment
CONCLUSIONS
91
93
REFERENCES
95
APPENDICES A
96
APPENDICES B
98
APPENDICES C
101
vi
LIST OF TABLES
TABLE NO.
TITLE
PAGE
1.1
Physical and Chemical Properties Propylene
1
1.2
Product and Application
3
1.3
Process Production of Propylene
4
1.4
Price of Propylene
11
1.5
Summary of Review and Screening of Alternative
13
Processes
1.6
Comparison between short-listed locations
16
2.1
Physical and Chemical Properties of Reactant
21
2.2
Physical and Chemical Properties of Product
21
2.3
Summary of Review and Screening of Alternative
24
Processes
2.4
Boiling Points of Propylene and Its Side Products
29
3.1
Mass Balance
47
3.2
Energy Balance
48
3.3
Stream Table Data
55
3.4
Shifted Temperatures
55
3.5
Summary of Temperature of Heat Exchanger
58
3.6
Heating Requirement Before and After Heat
59
Integration
3.7
Cooling Requirement Before and After Heat
59
Integration
4.1
Market Value of Propane, Hydrogen and Propene
62
6.1
Summary of Bare-module Cost for All Equipment
82
vii
LIST OF FIGURES
FIGURE NO.
TITLE
PAGE
1.1
Structural Formula of Propylene
2
1.2
Commercial Process Flow Diagram
6
1.3
Propylene Global Demand Profile
6
1.4
Global Propylene Sources Summary
7
1.5
Global Propylene Production by Process
8
1.6
Propylene Top Producers
9
1.7
Malaysia Polypropylene Demand, Exports,
10
Imports and Capacity
1.8
Global Propylene Prices
11
1.9
Typical product yields (%) by mass from steam
14
cracking various hydrocarbon feedstock
1.10
Pengurusan Air Pahang Berhad (PAIP)
17
1.11
Syarikat Air Johor
17
1.12
Syarikat Air Terengganu (SATU)
18
1.13
Tenaga Nasional Berhad (TNB)
18
2.1
Typical product yields (%) by mass from steam
24
cracking various hydrocarbon feedstock
2.2
Flowsheet with separation units of propylene
28
production process (Alternative 1)
2.3
Flowsheet with separation units of propylene
30
production process (Alternative 2)
2.4
Flowsheet with Temperature, Pressure and Phase
Change Operations in The Propylene Production
Process
31
viii
2.5
Flowsheet Task Integration for The Propylene
36
Production Process
3.1
Algorithm Table
56
3.2
Heat Exchanger Network
57
3.3
Temperature Profile H1
58
3.4
Temperature Profile H2
58
3.5
Temperature Profile H3
58
3.6
Temperature Profile H4
58
3.7
Process Flow Diagram Heat Exchanger Network
60
4.1
Solver Feature of Microsoft Excel
64
1
CHAPTER 1
PROJECT SELECTION
1.1
Background of Propylene
Propylene, also known as propene or methyl ethylene is a colourless and
flammable gases. It is an unsaturated organic compound which having chemical
formula of C3H6. It has one double bond which is the second simplest member of the
alkene class of hydrocarbons. The double bond presence in the propylene make it
boiling point is slightly lower than propane and thus more volatile. Propylene is comes
from cigarette smoke, combustion from forest fires, motor vehicle and air craft
exhaust. Propylene has low acute toxicity from inhalation and inhalation of this gas
can cause anaesthetic effects. Physical and chemical properties of propylene are shown
in Table 1.1.
Table 1.1: Physical and Chemical Properties Propylene [1]
Properties
Formula
Molecular weight (g/mol)
Value
C3H6
42.081
Boiling point (°C)
-48
Melting point (°C)
-185
Flash point (°C)
-108
2
Density (kg/m³ )
1.91
Solubility (mg/L)
Very soluble in water, 200
mg/L at 25°C
Colour
Colorless
Figure 1.1: Structural Formula of Propylene
Propylene is traded commercially in three grades which are chemical, polymer
and refinery. Chemical-grade propylene has minimum purity of 92-95%. Polymergrade propylene typically has minimum purity of 99.5-99.8% and contains the
impurities like propane, methane, ethane, ethylene, propyne, butenes, propadiene,
methylacetylene, butadiene, acetylene, diolefins, carbonyl sulfide, hydrogen, carbon
monoxide, carbon dioxide, oxygen, nitrogen, water and sulphur. Refinery-grade
propylene usually contain 50-70% propylene admixed with other low relative
molecular mass hydrocarbons.
1.2
Production and Uses of Propylene
Propylene is produced primarily as a by-product of petroleum refining and of
ethylene production by steam cracking of hydrocarbon feedstock. In steam cracking,
a mixed stream of hydrocarbons ranging from ethane to gas oils is pyrolysed with
steam. Product obtained in the process can be change to optimize production of
ethylene, propylene, or other alkenes by altering feedstock, temperature and other
parameters. The catalytic dehydrogenation of propane can also been used for the
production of propylene.
3
Propylene is a major industrial chemical intermediate that serves as one of the
building blocks for an array of chemical and plastic products and also the first
petrochemical employed in the industrial scale. The main uses of refinery propylene
are in liquefied petroleum gas (LPG) for thermal use or as an octane-enhancing
component in motor gasoline. The most important derivatives of chemical and polymer
grade propylene are polypropylene, propylene oxide, isopropanol, cumene and
acrylonitrile. Other commercial derivatives include acrylic acid and esters, oxo
alcohols and aldehydes, epichlorohydrin, synthetic glycerine and ethylene-propylene
copolymers. Table 1.2 below shows the main products of propylene and its application.
Table 1.2: Product and Application
1.3
Product
Application
Polypropylene
Mechanical parts, containers, fibres, films
Propylene oxide
Propylene glycol, antifreeze, polyurethane
Cumene
Polycarbonates, phenolic resins
Acrylonitrile
Acrylic fibres, ABS polymers
Oxo-alcohols
Coatings, plasticizers
Acrylic acid
Coatings, adhesives, super absorbent polymers
Propylene Manufacturing
Propylene is commercially generated as a co-product, either in an olefins plant
or a crude oil refinery’s fluid catalytic cracking (FCC) unit, or produced in an on
purpose reaction like propane dehydrogenation. The different process of the
production of propylene will be analysing before we choose the best reaction process
to produce propylene. The criteria of the selective reaction are based on their cost of
raw material, environmental impacts, safety, percentage yield of conversion and other
factors that will affect reaction process. Table 1.3 below shows several processes with
a lower cost propylene production.
4
Table 1.3: Process Production of Propylene
Process
Olefin Metathesis
Description

Metathesis is a reversible reaction between
ethylene and butene in which double bonds are
broken and then reform to form propylene.

Propylene yields about 90%.

This process may also been used when there is no
butene feedstock. For this case, part of the
ethylene will feeds an ethylene-dimerization unit
that convert ethylene into butene.
Propane

Dehydrogenation
A catalytic process that convert propane into
propylene and hydrogen (by product).

The yield of propylene is about 85%.

The reaction by product (mainly hydrogen) is
usually used as fuel for the propane
dehydrogenation reaction.

So, propylene tends to be the only product unless
local demand exists for the hydrogen by product.
Steam cracking

(naphtha)
In steam cracking, a gaseous or liquid
hydrocarbon feed like naphtha is diluted with
steam and briefly heated in a furnace without the
presence of oxygen.

The reaction temperature is very high, at around
850 °C, but the reaction is only allowed to take
place very briefly.

In modern cracking furnaces, the residence time is
reduced to milliseconds to improve yield.
5
1.4
Process Flow Diagram
The main objective of this process is to produce 100,000 Ib/hr of propylene with
polymer-grade propylene of 99.5% purity. Dehydrogenation of propane is chosen as
our desired reaction because of its simplest reaction. The commercial Process Flow
Diagram (PFD) is shown in the Figure 1.2. The plant consists of three main sections:
i.
Reactor
ii.
Catalyst Regeneration
iii.
Product recovery
Hydrocarbon feed is mixed with hydrogen-rich recycle gas and is introduced
into the heater to be heated into the desired temperature (over 540 °C) and then enter
the reactors to be converted at high mono-olefin selectivity. Several interstage heaters
are used to maintain the conversion through supplying heat continuously since the
reaction is endothermic.
Catalyst activity is maintained by continuous catalyst regenerator (CCR) or
shutting down reactors one by one and regenerating the reactor by the regeneration air,
the continuous catalyst regenerator is where the catalyst is continuously withdrawn
from the reactor, then regenerated, and fed back to the reactor bed. Reactor effluent is
compressed, dried and sent to a cryogenic separator where net hydrogen is recovered.
The olefin product is sent to a selective hydrogenation process where dienes and
acetylenes are removed. The propylene stream goes to a de-ethanizer where light-ends
are removed prior to the propane-propylene splitter. Unconverted feedstock is recycled
back to the depropanizer where it combines with fresh feed before being sent back to
the reactor section.
6
Figure 1.2: Commercial Process Flow Diagram
1.5
Market Survey
1.5.1
Introduction
Propylene is perhaps the most versatile building block in the petrochemical
industry, in terms of its variety of end-use products and its multitude of production
sources. In 2015, the global demand for propylene, polymer grade and chemical grade
combined, is estimated at 94.2 million tonnes. The chart in Figure 1.3 outlines the
propylene global demand profile for 2015.
Figure 1.3: Propylene Global Demand Profile
7
1.5.2
Production of Propylene
Propene production increased in (Europe and North America only) from 2000
to 2008, it has been increasing also in East Asia, most notably Singapore and
China. Total world production of propene is currently about half that of ethylene.
About 56% of the worldwide production of propylene is obtained as a co-product of
ethylene manufacture, and about 33% is produced as a by-product of petroleum
refining. About 7% of propylene produced worldwide is on-purpose product from the
dehydrogenation of propane and metathesis of ethylene and butylenes; the remainder
is from selected gas streams from coal-to-oil processes and from deep catalytic
cracking of vacuum gas oil (VGO). The supply of propylene remains highly dependent
on the health of the ethylene industry as well as on refinery plant economics.
The chart in Figure 1.4 summarizes the global propylene sources in 2015.
Figure 1.4: Global Propylene Sources Summary
Since lighter feedstock is used, it has reduced propylene co-product production
resulting in an increased investment in on-purpose production. As shown in Figure 6,
8
30% of global propylene supply will be from unconventional sources by 2025. PDH
has been through a phase of major expansion in the Middle East but the focus is now
moving to China and the US.
Figure 1.5: Global Propylene Production by Process
1.5.3
Propylene Consumption
The market dynamic of propylene is influenced by polypropylene.
Polypropylene accounts for around 65% of global propylene production, ranging from
53% in North America to more than 90% in Africa and the Middle East.
Top world companies are leading the production of propylene with
LyondellBasell, Netherlands on top of propylene producing companies by 2009. The
top propylene-producing companies are listed as bellow:
9
Propylene Top Producers
Relliance Industries, India
Formosa Plastics Group, Taiwan
PetroChina, Bejing, China
SABIC, KSA
ExxonMobil Chemical, USA
Ineos Group, England
Total SA, France
Sinopec, China
LyondellBasel Industries, Netherlands
3.8
4
4
4
5
5.8
6.1
9
15.8
Figure 1.6: Propylene top producers
As shown in Figure 1.6, after experiencing zero growth or declines in 2008 and
2009, global propylene consumption grew at a rate of almost 7.5% in 2010, led by
Asia at 11% year-on-year. The economic recession of 2008/2009 reflected both a
reduction in pull-through demand for polypropylene, as well as a supply-chain
inventory rundown, reminiscent of the early 1980s downturn. World petrochemical
industries have historically witnessed very few upheavals that combined the effects of
both energy volatility and depressed downstream demand.
The fifteen largest worldwide producers of propylene accounted for almost
51% of world capacity as of 2010, representing about the same level of concentration
as five years ago. The most significant changes in the last two years have been Sinopec
taking over the top spot, a position long occupied by ExxonMobil, and PetroChina
jumping from the seventh spot to number four.
World consumption of propylene is forecast to grow slightly better than global
gross domestic product (GDP) rates over the next five years. Average growth will be
5% per year, higher than GDP in general and higher than ethylene specifically, with
growth for polypropylene being much better than that for polyethylene. Growth will
10
be led by the Middle East, Asia, Central and Eastern Europe, and South America at
12.5%, 6.5%, 5%, and 4.5% per year, respectively. Asia is a mixed bag of growth rates
with China and India at 8–10% annually and the mature economies of Japan, the
Republic of Korea, and Taiwan at 1–2% per year. Near-term growth will be relatively
slow in the mature economies of North America and Western Europe.
1.5.4
Local outlook for production of propylene
Malaysia’s petrochemical sector has contributed significantly to the
development of local downstream plastic processing activities. Malaysia is one of the
largest plastics producers in Asia, providing a steady supply of feedstock materials for
the plastic processing industry such as propylene.
Figure 1.7: Malaysia Polypropylene Demand, Exports, Imports and Capacity
11
1.5.5
Market price of propylene
Propylene production from crackers have been constrained due to low
ethylene prices, which prompted an increase in lighter feedstocks like ethane,
which produces the least amount of co-products like propylene. With ethane
feedstock
costs higher and volatile,
ethylene
spot
prices
have
risen
and naphtha cracker margins are no longer negative. But the demand fall-off has
not
been
as
steep
as
expected
amid
a
tight
market
for
downstream polypropylene (PP). While imports of PP have increased, domestic
operating rates have remained relatively good. PP is the largest consumer of US
propylene, but indications for other downstream sectors are similar. A market
source said that demand seems to be strong across most major derivatives.
Figure 1.8: Global Propylene Prices
Table 1.4: Price of Propylene
Price
Product
Propylene
RM/Ib
RM/Kg
1.49
3.285
12
1.6
Process Screening
1.6.1
Gross Profit
There are three methods of producing propylene in industry. Each method uses
different raw material and vary according to price. The methods are:
1. Olefin Metathesis
2. Propane Dehydrogenation
3. Steam cracking
All the three methods will be screened and will be chosen based on the criteria on
economic potential. Conversion rate used for the following calculations is 1 USD =
RM 4.14.
Gross profit will be calculated based on this formula:
𝐺𝑟𝑜𝑠𝑠 𝑝𝑟𝑜𝑓𝑖𝑡 = 𝑃𝑟𝑜𝑑𝑢𝑐𝑡 𝑣𝑎𝑙𝑢𝑒 − 𝑅𝑎𝑤 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 𝑣𝑎𝑙𝑢𝑒
The stoichiometry of the reactant and products are taking into account without
considering the side products.
1) Olefin Metathesis is a reversible reaction between ethylene and butylene in which
double bonds are broken and then reform to form propylene.
𝐶2 𝐻4 + 𝐶4 𝐻8 → 2𝐶3 𝐻6
Chemical
Ethylene
Butylene
Propylene
C2H4
C4H6
C3H6
1
1
2
28.05
56.10
42.08
Mass
28.05
56.10
84.16
Kg/ Kg Propylene
0.3333
0.6666
1
1.39
0.65
1.05
Molecular formula
Kgmol
Molar
mass
(kg/kgmol)
USD/Kg
13
RM/Kg
5.75
2.62
4.36
Gross Profit = 4.36(1) - 5.75(0.3333) – 2.62(.6666) = RM 0.70/ Kg Propylene
2) Propane Dehydrogenation is a catalytic process that convert propane into
propylene and hydrogen
𝐶3 𝐻8 → 𝐶3 𝐻6 + 𝐻2
Chemical
Propane
Molecular formula
Hydrogen
Propylene
C3H8
H2
C3H6
1
1
2
44.09
2.016
42.08
Mass
44.09
2.016
42.08
Kg/ Kg Propylene
1.0478
0.048
1
USD/Kg
0.43
0
1.05
RM/Kg
1.76
0
4.36
Kgmol
Molar
mass
(kg/kgmol)
Gross Profit = 4.36(1) – 1.76(1.0478) – 0(0.048) = RM 2.52/ Kg Propylene
3) Steam Cracking is where saturated hydrocarbons are broken down into smaller,
often unsaturated, hydrocarbons.
However, propylene is merely a by-product, it is synthesized by other methods,
such as propane dehydrogenation. Ethylene is mainly produced in steam cracking.
Figure 1.9 shows the typical product yields (%) by mass from steam cracking various
hydrocarbon feedstock.
14
Figure 1.9: Typical product yields (%) by mass from steam cracking various
hydrocarbon feedstock
As it can be seen from the figure, it is not feasible to consider using steam cracking to
produce propylene as the highest yield (%) is 19.
1.6.2
Factors Affecting Screening Process
Table 1.5: Summary of Review and Screening of Alternative Processes
Olefin Metathesis
C2H4 + C4H8  2C3H6
Gross Profit
RM 0.70/Kg propylene
Dehydrogenation of
propane
C3H8  C3H6 + H2
RM 2.52/Kg propylene
Butane and ethylene is
Safety
flammable, and ethylene also
Propane is flammable.
may cause dizziness
Catalyst
Al2O3
Al2O3
By-product
No by-product
Hydrogen
Operating
Temperature: 90-100ºC
condition
Pressure: 100 – 110 bar
Conversion
90% percent of conversion
86% percent of conversion
Flammability
Flammable
Flammable
Temperature: 560 – 650 ºC
Pressure : slightly below
atmospheric pressure
1.6.2.1 Temperature
The operating temperature of olefin metathesis process is lower than the
dehydrogenation of propane. High temperature process unit are not economically
feasible because it takes large amount of energy to reach the temperature. Thus, olefin
15
metathesis makes the case of being more economically feasible and safer to work as
well as high temperature carries risk.
1.6.2.2 Pressure
Dehydrogenation of propane operates at nearly atmospheric pressure, which is
more economically feasible as it is a well-known fact that compressing units are
expensive. Which is why even if olefin metathesis operates at lower temperature, it is
still more expensive to bring the pressure to 100 bars.
1.6.2.3 Safety
The aspect of safety carries a risk, as both methods are flammable. The catalyst
used as well is same and Al2O3 does not carry any serious health apart causing irritation
to eyes and skin upon contact.
1.6.2.4 Environment
None of the chemicals used carries any exotoxicity. Thus, the two methods
does not pose a major threat to the environment.
In conclusion, dehydrogenation of propane is better economic choice as the
difference between two methods gross profit is RM 1.82/ Kg propylene. Furthermore,
the operating cause would be slightly cheaper for dehydrogenation of propane due to
the low-pressure requirement compared to olefin metathesis. Hence, dehydrogenation
of propane of propane is chosen.
16
1.7
Site Location
Location for the chemical plant plays very important role because it can affect
the plant operation and its success. For the construction of the propylene plant, we
have listed three main industrial estates that are located at south and east coast of
Peninsular Malaysia. The industrial lands are

Tanjung Langsat Industrial Land Pasir Gudang, Johor

Gebeng Industrial Land,Pahang

Kerteh Industrial Land, Terengganu
Table 1.6: Comparison between short-listed locations
Specification
Location
Land Availability
(Hectare)
Land Price (per
square feet)
Raw
Material
Supply
Tanjung Langsat, Gebeng, Pahang
Kerteh,
Johor
Terengganu
Pasir
Gudang Gebeng Industrial Kerteh Industrial
Industrial Land
Land
Land
4
250
4.98
RM 71
Rm 15
RM 30
-Titan
Petrochemicals
Sdn Bhd
-Peninsular
Gas
Utilisation Project
– Gas Malaysia
Berhad
-North-South
Transport
Expressway
Facilities
(PLUS)
-Senai
International
Airport
-Tanjung Langsat
Port Facilities
Port
-Johor Port
-Tanjung Pelepas
Port
Distance
from 42km from Johor
Bahru
town
Tenaga Nasional
Power Supply
Berhad (TNB)
Petronas
Berhad
Gas -Petronas
Gas
Berhad
-Petronas
Penapisan
(Terengganu) Sdn
Bhd
-Kuantan-Kerteh
Railway
-East
Coast
Expressway
-East
Coast
Expressway
-Kuantan-Kerteh
Railway
-Kerteh Airport
Kuantan Port City -Kertih Terminal
(KPC)
-Kertih Port
40km
from
Kuantan
Tenaga Nasional
Berhad (TNB)
42km
from
Kemaman
Tenaga Nasional
Berhad (TNB)
17
Syarikat Air Johor -Semambu
(SAJ)
Reservoir
-Pengurusan Air
Pahang
Berhad
(PAIP)
Chemical
and
Chemicals
and
Type of Preferred
Petrochemical
Petrochemical
Industry
-Universiti
Availability
of -Industrial
Training
Malaysia
Labor
Institute
(ITT) Pahang
Johor
-Pusat
-Universiti
Pembangunan
Teknologi
Kemahiran
Malaysia (UTM)
Pahang
Water Supply
Kualiti Alam
Environmental
Effect &Effluent
Disposal
Hot and Humid
Climate
1.7.1
Water Tariff
-Syarikat
Terengganu
(SATU)
Air
Kualiti Alam
Chemical
and
Petrochemical
-Pusat
Pembangunan
Kemahiran Negeri
Terengganu
(TESDEC)
-Universiti
Malaysia
Terengganu
(UMT)
Kualiti Alam
Hot and Humid
Hot and Humid
18
Figure 1.10: Pengurusan Air Pahang Berhad (PAIP)
Figure 1.11: Syarikat Air Johor (SAJ)
Figure 1.12: Syarikat Air Terengganu (SATU)
Based on Table 2,3 and 4, the price rate is RM RM 0.84/m 3, RM 3.30/m3 and
RM 1.15/m3 respectively. Therefore, the cheapest price for PAIP (Pahang) which is
RM 0.84/m3 while SAJ (Johor) has the highest price which is RM 3.30/m 3.
19
1.7.2
Electricity Tariff
Figure 1.13: Tenaga Nasional Berhad (TNB)
The industrial price of tariff rate for medium voltage general industrial tariff in
TNB is RM RM 37/kW, for each kilowatt of maximum demand per month. The price
rate is the same for all locations since TNB is the sole electricity supplier in Peninsular
Malaysia.
1.8
Conclusion
After considering all the short-listed locations, we have decided to build the
proposed 100,000 MTA propylene plant at Gebeng Industrial Land, Pahang. This is
due to the most important factor which is the price of the land is the cheapest in
Gebeng, RM 15/psf and has the largest area of land compared with Tanjung Langsat
and Kertih. Next, gebeng is also near to the East Coast Expressway and KuantanKerteh Railway which made our land transportation easier for raw material
transportation from Petronas Gas Berhad. For export or import purposes, Gebeng is
also near to the Kuantan Port City. The water tariff in Pahang is also the cheapest in
Pahang compared with Johor and Terengganu. In conclusion, we choose Gebeng due
to cheaper utilities costs and its strategic location.
20
CHAPTER 2
PROCESS CREATION AND SYNTHESIS
2.1
Introduction
By performing process screening based on gross profit (GP) and other factors
related to sustainable design for all possible reaction pathways, the dehydrogenation
of propane is chosen as the best reaction pathway to produce 100,000 lb/hr of
propylene. Next, Gebeng Industrial Land, Pahang is selected as the location for the
establishment of propylene manufacturing plant compared to Tanjung Langsat
Industrial Land, Johor and Kertih Industrial Land, Terengganu due to its lower land
price and adequate utilities.
2.2
Sources of Raw Material and Its Specification
The raw material needed to produce propylene is propane with catalyst,
aluminum oxide. Propane is suggested to be bought from Petronas Gas Berhad
Malaysia due to its near location which is located in Semambu Industrial Land,
Kuantan. As for the aluminium oxide, it will be brought from Superb Aluminum
Industries,Selangor. By using the available East Coast Expressway, the time needed
for transportation is quite short which about 3 hours is.
21
2.3
Physical and Chemical Properties
Table 2.1: Physical and Chemical Properties of Reactant
Properties
Propane
Chromium (catalyst)
C3H8
Al2O3
44.096
101.96
Boiling point (°C)
-42
3000
Melting point (°C)
-189.7
2030
Density (kg/m³ )
2.01 at 0°C
3970 at 0°C
Solubility (mg/L)
Very soluble in water,
Insoluble in water
Formula
Molecular weight (g/mol)
0.0624 mg/Ml at 25°C
Appearance
Colourless
White crystalline
powder
Odor
Std
enthalpy
of
formation ΔfHo298 (Kj/mol)
Odorless
Odorless
-119.8 (l)
-1675.7
-103.8 (g)
Table 2.2: Physical and Chemical Properties of Product
Properties
Propylene
Hydrogen (side
product)
Formula
C3H6
H2
42.081
2.016
Boiling point (°C)
-48
-253
Melting point (°C)
-185
-259.2
Density (kg/m³ )
1.81 at 15°C
0.08988 at 0°C
Solubility (mg/L)
Very soluble in water,
0.00162 mg/mL at 21°C
Molecular weight (g/mol)
0.2 mg/mL at 25°C
Appearance
Odor
Std
enthalpy
of
formation ΔfHo298 (kJ/mol)
Colorless
Colorless
Practically odorless
Odorless
+20.41
0
22
2.4
Synthesis Steps
Process synthesis involves the selection of processing operations to convert
raw materials to products, given that the states of the raw material and product streams
are specified. The most widely accepted approach for process synthesis is introduced
by Rudd, Powers, and Siirola (1973) in a book entitled Process Synthesis. There are 5
key synthesis steps which are:
1. Eliminate differences in molecular types
2. Distribute the chemicals by matching sources and sinks
3. Eliminate differences in composition
4. Eliminate differences in temperature, pressure, and phase
5. Task integration; combination of operations into unit processes and decide
between continuous and batch processing
Step 1 - Eliminate Differences in Molecular Type
There are three methods of producing propylene in industry. Each method uses
different raw material and vary according to price. The methods are:
1. Olefin Metathesis
2. Propane Dehydrogenation
3. Steam cracking
All the three methods will be screened and will be chosen based on the criteria on
economic potential. Conversion rate used for the following calculations is 1 USD =
RM 4.14.
Gross profit will be calculated based on this formula:
𝐺𝑟𝑜𝑠𝑠 𝑝𝑟𝑜𝑓𝑖𝑡 = 𝑃𝑟𝑜𝑑𝑢𝑐𝑡 𝑣𝑎𝑙𝑢𝑒 − 𝑅𝑎𝑤 𝑚𝑎𝑡𝑒𝑟𝑖𝑎𝑙 𝑣𝑎𝑙𝑢𝑒
23
The stoichiometry of the reactant and products are taking into account without
considering the side products.
1) Olefin Metathesis is a reversible reaction between ethylene and butylene in which
double bonds are broken and then reform to form propylene.
𝐶2 𝐻4 + 𝐶4 𝐻8 → 2𝐶3 𝐻6
Chemical
Ethylene
Butylene
Propylene
C2H4
C4H6
C3H6
1
1
2
28.05
56.10
42.08
Mass
28.05
56.10
84.16
Kg/ Kg Propylene
0.3333
0.6666
1
USD/Kg
1.39
0.65
1.05
RM/Kg
5.75
2.62
4.36
Molecular formula
Kgmol
Molar
mass
(kg/kgmol)
Gross Profit = 4.36(1) - 5.75(0.3333) – 2.62(.6666) = RM 0.70/ Kg Propylene
2) Propane Dehydrogenation is a catalytic process that convert propane into
propylene and hydrogen
𝐶3 𝐻8 → 𝐶3 𝐻6 + 𝐻2
Chemical
Propane
Molecular formula
Hydrogen
Propylene
C3H8
H2
C3H6
1
1
2
44.09
2.016
42.08
Mass
44.09
2.016
42.08
Kg/ Kg Propylene
1.0478
0.048
1
USD/Kg
0.43
0
1.05
RM/Kg
1.76
0
4.36
Kgmol
Molar
mass
(kg/kgmol)
24
Gross Profit = 4.36(1) – 1.76(1.0478) – 0(0.048) = RM 2.52/ Kg Propylene
3) Steam Cracking is where saturated hydrocarbons are broken down into smaller,
often unsaturated, hydrocarbons.
However, propylene is merely a by-product, it is synthesized by other methods,
such as propane dehydrogenation. Ethylene is mainly produced in steam cracking.
Figure 1.9 shows the typical product yields (%) by mass from steam cracking various
hydrocarbon feedstock.
Figure 2.1: Typical product yields (%) by mass from steam cracking various
hydrocarbon feedstock
As it can be seen from the figure, it is not feasible to consider using steam cracking to
produce propylene as the highest yield (%) is 19.
Table 2.3: Summary of Review and Screening of Alternative Processes
Olefin Metathesis
C2H4 + C4H8  2C3H6
Gross Profit
RM 0.70/Kg propylene
Dehydrogenation of
propane
C3H8  C3H6 + H2
RM 2.52 / Kg propylene
25
Butane and ethylene is
Safety
flammable, and ethylene also
Propane is flammable.
may cause dizziness
Catalyst
Al2O3
Al2O3
By-product
No by-product
Hydrogen
Operating
Temperature: 90-100ºC
condition
Pressure: 100 – 110 bar
Conversion
90% percent of conversion
86% percent of conversion
Flammability
Flammable
Flammable
Temperature: 560 – 650 ºC
Pressure : slightly below
atmospheric pressure
In conclusion, dehydrogenation of propane is better economic choice as the
difference between two methods gross profit is RM 1.82/ Kg propylene. Furthermore,
the operating cause would be slightly cheaper for dehydrogenation of propane due to
the low-pressure requirement compared to olefin metathesis. Hence, dehydrogenation
of propane of propane is chosen. The operating temperature of olefin metathesis
process is lower than the dehydrogenation of propane. High temperature process unit
are not economically feasible because it takes large amount of energy to reach the
temperature. Thus, olefin metathesis makes the case of being more economically
feasible and safer to work as well as high temperature carries risk. Dehydrogenation
of propane operates at nearly atmospheric pressure, which is more economically
feasible as it is a well-known fact that compressing units are expensive. Which is why
even if olefin metathesis operates at lower temperature, it is still more expensive to
bring the pressure to 100 bars. The aspect of safety carries a risk, as both methods are
flammable. The catalyst used as well is same and Al 2O3 does not carry any serious
health apart causing irritation to eyes and skin upon contact. None of the chemicals
used carries any ecotoxicity. Thus, the two methods do not pose a major threat to the
environment.
26
Step 2 – Distribute the Chemicals
Reactor
T = 600 OC
m1C3H8
F lb/hr C3H8
P = 1 bar
m2C3H6
m3H2
R lb/hr C3H8
Overall Reaction Equation:
C3H8
C3H6 + H2
Basis: 100000 lb/hr of propylene (C3H3)
86% of conversion
C3H8
C3H6
H2
1
1
1
m1
m2=100,000.00
m3
MW (lb/lbmol)
44.10
42.08
2.01
n, (lbmole/hr)
2376.43
2376.43
2376.43
stoichiometry
Mass flowrate (lb/hr)
27
Number of moles of propylene formed
= (100,000 lb/hr)/42.08
= 2376.43 lbmole/hr C3H6
Assume 100% conversion, the mass flow rate of feed, m1 = 2376.43 x 44.1
= 104800.56 lb/hr
for 86% conversion, the mass flow rate of recycle, R = (1-0.86)/0.86 x 104800.56
R = 17060.53 lb/hr
Mass flowrate of H2 , m3
= (no. of mole) X (molecular weight)
= (2376.43) x (2.01)
= 4776.62 lb/hr
Mass flowrate feed to the reactor,
F = m1 + R
= 104800.56 + 17060.53 = 121861.09l
28
Step 3: Eliminate Differences in Composition
Alternative 1:
Figure 2.2: Flowsheet with separation units of propylene production process
(Alternative 1)
In the production of propylene, one reactor and three separators are used in
order to enable all chemicals involved to be supplied to their sinks. Figure 2.1 shows
the separation units that are needed in a propylene production process. The raw
material for the production of propylene consists of 0.98 propane and 0.02 ethane.
Even though the feed is not completely pure, but no separating unit is needed because
there is only small proportion of ethane in propane. The feed will then enter the reactor
at 600 °C and 1 atm. These pressure and temperature is selected because the
dehydrogenation process of propylene only will occur at these conditions [2].
After the reaction occurs, there are a lot of products produced from the
dehydrogenation process. In order to separate the products, 3 separation units will be
used. The first product that will be separated is hydrogen gas. The reason is, hydrogen
29
gas has a low value of critical pressure at 12.96 bar and it will be difficult to separate
the other products if the hydrogen maintain in the product mixtures. S2 will be used as
separation unit that will be operated at pressure 1 atm and at cryogenic temperature of
-129 °C [3] since hydrogen gas has low critical temperature at -240.01 °C. At column
temperature of -129 °C, hydrogen gas will leave the column as a vapour at the distillate
whereas propylene, propane and other side products will leave the column as liquid at
bottom.
Next, after separate hydrogen gas, we will separate propane and propylene
from the side product. From Table 2.4 at 1 atm, the boiling point of C 3 is very low, 47.7 ˚C, and hence if C3 were recovered at 1 atm as the distillate of the S3, very costly
refrigeration would be necessary to condense the reflux stream. At 17.5 bar, the bubble
point of propane and propylene mixture is at 40 ˚C. By operating the distillation
column at 20 ˚C, the propane and propylene mixture will leave the column as liquid at
the bottom.
The last unit operation would be used to separate the bottom products of second
separator into nearly pure species which is specified at 15 bar. Under these conditions,
the distillate (nearly pure propylene) boils at 38 ˚C. By operating the column at 40 ˚C,
the propylene will leave as the distillate and can be condensed with inexpensive
cooling water, which is available at 25˚C whereas the unreacted propane will leave at
the bottom and recycled back to the reactor.
Table 2.4: Boiling Points of Propylene and Its Side Products
Chemical
H2
C3H8
C3H6
C2H6
C2H4
Normal boiling
point
(1atm, ˚C)
-252.78
-42.00
-47.7
-89.00
-103.7
Boiling point (˚C)
15 bar
46.54
33.06
-16.02
-37.22
17.5 bar
53.62
44.95
-10.13
-31.84
30
Alternative 2:
Figure 2.3: Flowsheet with separation units of propylene production process
(Alternative 2)
At 17.5 bar, the bubble point of propane and propylene mixture is at 40˚C while
ethane and ethylene mixture is at -21˚C. By operating the distillation column (S2) at
20˚C, the propane and propylene mixture will leave the column as liquid at the bottom
while ethane and ethylene mixture will leave the column as the distillate.
The distillation column (S3) would be used to separate the upper product from
S2 into nearly pure species which is specified at 41 bar. Under these conditions, the
distillate (nearly pure ethylene) boils at 2˚C. By operating the column at 10˚C, the
ethylene will leave as the distillate whereas the ethane will leave at the bottom.
Since our main product is propylene, it is not necessary to separate the ethane
and ethylene mixture. This is because separation of ethane and ethylene mixture will
increase the cost as we need to build another distillation column to separate it.
Therefore, alternative 1 is chosen for the production of propylene.
31
Step 4: Eliminate Differences in Temperature, Pressure, and Phase
Figure 2.4: Flowsheet with Temperature, Pressure and Phase Change Operations in The Propylene Production Process
32
Figure 2.3 shows the changes of the state of chemicals. Since the original state of the
raw material is at 36 °C and 11 bar, its temperature is raised to 600 °C at 1 bar.
The process begins by mixing the raw materials (ethane and propane gas) with a
stream of recycle propane gas at 36 °C and 11 bar. The mixing of raw materials and
recycle propane undergoes the following operations:
1. The mixture is preheated before it is introduced to the reactor. The reaction
occurs at around 600oC and 1 bar.
2. The product mixture is then cooled to its dew point -129oC at 1 bar.
3. Then, the product mixture is introduced into a condenser (S1) that separates
the hydrogen gas from other liquid products.
In addition, the liquid mixture that condensed at -129oC at 1 bar from the condenser
is operated upon as follows:
1. Its pressure is increased to 17.5 bar.
2. The temperature is then raised to a liquid at its bubble point, 20 oC at 17.5 bar.
3. Then, the liquid mixture is introduced into a separation column (S2) that
separates the propane gas and propylene gas from other liquid products.
Next, the bottom products (propane gas and propylene gas) from separation
column (S2) are then entered into separation column (S3) at 40 oC. The propylene gas
with a boiling point of 33oC at 15 bar is come out as an upper product from separation
column (S3). The upper products (ethane gas and ethylene gas) from separation
column (S2) will be sell off to market. Finally, the propane liquid from the recycle
stream (at 40oC and 15 bar) undergoes the operation where its temperature is lowered
to the mixing temperature at 36oC at 11 bar.
33
Step 5: Task Integration
Figure 2.5 below shows task integration for the process of propylene production.
Task integration is where the selection of processing units, often referred as unit
operations, in which one or more of the basic operations are carried out. At this stage
in process synthesis, it is common to make the most obvious combinations of
operations, leaving many possibilities to be considered when the flowsheet is
sufficiently promising to undertake the preparation of a base-case design. Below are
the descriptions of unit processes shown in Figure 2.4:
1. Control valve
Control valve is used to regulate pressure and available for any pressure. In this
process, it is used to reduce the high pressure of the fluid to a desired pressure
which is 1 bar. The first control valve is used to reduce the pressure of the gas
in fresh feed from 11 bar to 1 bar. Another control valve is used to reduce the
pressure of liquid propane that is used to recycle back into the feed.
2. Mixer
The used of mixer is to mix gas from the fresh feed and recycle propane from
the splitter at desired temperature and pressure.
3. Furnace
An industrial furnace or direct fired heater is equipment used to provide heat
for a process or can serve as reactor which provides heat of reaction. Furnace
designs vary as to its function, heating duty, type of fuel and method of
introducing combustion air. Since the reactor that we used required high
temperature which is 600˚C and the outlet temperature from the mixer is low
at 36°C, the furnace is used to heat up the stream. This follows heuristics 25
which explained unless required as part of the design of the separator or reactor,
provide necessary heat exchange for heating or cooling process fluid streams,
with or without utilities, in an external shell-and-tube heat exchanger using
countercurrent flow. However, if a process stream requires heating above
750°F (400°C), use a furnace unless the process fluid is subject to chemical
decomposition.
34
4. Oleflex Reactor
This type of reactor involves a process of catalytic dehydrogenation for the
production of light olefins from their corresponding paraffin. One of the
processes of catalytic dehydrogenation is production of propylene from
propane. The Oleflex process provides high quality of propylene, which then
leads to high quality polymers. This process used Pt-Sn catalyst to promote the
dehydrogenation reaction. In this process, the Oleflex reactor operated at
600˚C and 1 bar.
5. Heat Exchanger
A heat exchanger is a device designed to efficiently transfer or "exchange" heat
between two or more fluids. In other words, heat exchanger is used in both
heating and cooling processes. The fluids may be separated by a solid wall to
prevent mixing or they may be in direct contact. In this process, heat exchanger
is needed to increase or decrease the temperature of the stream.
6. Pump
The used of pump is to provide sufficient pressure to overcome the operating
pressure
of
the
system
to
move
fluid
like
liquid at a
required
flowrate. To achieve a required flow through a pumping system, we need to
calculate what the operating pressure of the system will be to select a suitable
pump. Since the pressure change operation involves in this process is liquid, it
is accomplished by a pump. The enthalpy change in the pump is very small and
the temperature does not change by more than 1˚C which means that used of
pump are not affect temperature of the stream.
7. Cryogenic Separator (S1)
The required temperature and pressure for the separation to occur is 129˚C and
1 bar. The used of this separator is to remove hydrogen from the mixture
vapour. Hydrogen is removed at the distillate while other components are
removed at the bottom as the liquid.
35
8. Distillation column (S2)
Distillation is a process where a liquid or vapor mixture of two or more
substances is separated into its component fractions of desired purity, by the
application and removal of heat. The process is based on the fact that the vapor
of a boiling mixture will be richer in the components that have lower boiling
points. Therefore, when this vapor is cooled and condensed, the condensate
will contain more volatile components. At the same time, the original mixture
will contain more of the less volatile material. The equipment used for
distillation process is distillation column. In this process, the column is separate
the components of C3 and C2. Components of C2 are removed at the distillate
since their boiling points are lower than C 3 components while C3 components
are removed at the bottom as a liquid.
9. Splitter (S3)
Since propane and propylene have similar molecular size and physical
properties, their separation is challenging. Therefore, propane-propylene
splitter is used as it can give high purity of propylene. The purified propylene
recovered at the top is condensed, and the other part of it is returned as reflux
while the propane is drawn off at the reboiler.
36
36°C
11 bar
1
C3H8
C2H6
36 °C
1 bar
36 °C
1 bar
3
2
V-100
Mix-100
600
°C
1 bar
H2
5
C2H4
C3H6
C3H8
Furnace
R-100
6
C2H6
600 °C
1 bar
Condenser duty
E-100
-103.7°C
1 bar
H2
7
C3H6
C2H4
C2H6
14
19
9
Rectifier
4
-129°C
1 bar
T-100
20 °C
17.5 bar
E-101
8
17.5
bar
13
E-103
18
T-101
V-100
V-101
V-102
15 bar
40 °C
15 bar
E-105
12
1 bar
17
-10.13°C
17.5 bar
38.1°C
15 bar
T-100
T-101
S-100
C3H6
C2H4
10
C2H6
15
E-102
C3H8
C3H8
C3H6
20 °C
17.5 bar
11 -129°C
17.5 bar
-129°C
1 bar
C3H8
20
Reboiler duty
E-104
40 °C
15 bar
16
21
20 °C
15 bar
P-100
V-101
22
36 °C
1 bar
E-106
40 °C
1 bar
Labelling
E-100
E-101
E-102
E-103
E-104
E-105
E-106
Mix-100
P-100
R-100
S-100
V-102
40 °C
15 bar
Figure 2.5: Flowsheet Task Integration for The Propylene Production Process
Description
Heater
Heat Exchanger
Evaporator
Heater
Evaporator
Heater
Cooler
Mixer
Pump
Oleflex Reactor
Cryogenic
Separator
Distillation
Column
Splitter
Control valve
Control valve
Control valve
37
2.5
Manual Mass Balance Calculation
2.5.1
Overall Mass Balance
Reactor
T = 500 OC
m1C3H8
F C3H8
100000
C3H6
P = 1 bar
gfrgfr
R C3H8
Overall Reaction Equation :
C3H8
C3H6 + H2
Basis : 100000 lb/hr of propylene (C3H3)
86% of conversion
C3H8
H2
C3H6
1
1
1
m1
m2
100,000.00
MW (lb/lbmol)
44.10
2.01
42.08
n, (lbmole/hr)
2376.43
2376.43
2376.43
stoichiometry
Mass flowrate (lb/hr)
n C3H8 = 100 000/42.08
= 2376.43 lbmole/hr
m2 = 2376.43 × 2.01
= 4776.62 lb/hr
m1 = 2376.43 × 44.10
= 104 800.56 lb/hr
R=
1−0.86
0.86
× 104 800.56
= 17 060.56 lb/hr
F = 104 800.56 + 17 060.56
= 121 861.12 lb/hr
38
2.5.2
Mass Balance for Mixer
F1
M-1
0.98 C3H8
F2
C3H8
0.02 C2H6
C2H6
1
2
9
R = 17060.53 lb/hr
C3H8
Stream 1
Stream 2
Stream 9
Mole
Fraction
Mass
Flowrate,
(lb/hr)
104
800.56
0.983
121861.09 1.0
17060.5
3
2029.78
0.017
2029.78
0
No.
Component
Mole
Fractio
n
Mass
Flowrate,
(lb/hr)
1
Propane
0.98
2
Ethane
0.02
Mole
Fraction
0
Mass
Flowrate
, (lb/hr)
39
2.5.3
Mass Balance for Reactor
2
3
R-1
123 999.88 lb/hr
123999.88 lb/hr
3
C3H8
0.98 C3H8
C4H10
0.017 C2H6
C5H12
C3H6
C4H8
H2
The percentage of conversion for propane and ethane are 86% and 60% respectively.
Stream 2
Stream 3
No.
Component
Mole
Fraction
Mass Flowrate,
(lb/hr)
1
Propane
0.983
121 861.09
0.0737
17 060.53
2
Ethane
0.017
2 029.78
0.0052
811.912
3
Propylene
0
0
0.4530
100 000
4
Ethylene
0
0
0.0077
1137.99
5
Hydrogen
0
0
0.4604
4858.15
Mole
Fraction
Mass Flowrate,
(lb/hr)
40
2.5.4
Mass Balance for Separation Unit 1
4
H2
3
S-1
123 868.582 lb/hr
C3H8
C2H6
C3H6
5
C2H4
B1 lb/hr
H2
C3H8
C2H6
C3H6
C2H4
Stream 3
Stream 4
Stream 5
No
.
Componen
t
1
Propane
0.0737
17 060.53
-
-
0.1367
17
060.53
2
Ethane
0.0052
811.91
-
-
0.0096
811.91
3
Propylene
0.4530
100 000
-
-
0.8395
100 000
4
Ethylene
0.0097
1137.99
-
-
0.0143
1137.99
5
Hydrogen
0.4604
4858.15
1
4858.15
-
-
Mole
Fraction
Mass
Flowrate,
(lb/hr)
Mass
Flowrate,
(lb/hr)
Mole
Fraction
Mole
Fraction
Mass
Flowrate
, (lb/hr)
41
2.5.5
Mass Balance for Separation Unit 2
6
C2H6
C2H4
5
S-1
119 010.43 lb/hr
C3H8
C2H6
C3H6
7
C2H4
C2H6
C3H6
Stream 5
Stream 6
Stream 7
No
.
Component
Mole
Fraction
1
Propane
0.1367
17 060.53
-
-
0.14
17
060.53
2
Ethane
0.0096
811.91
0.4
811.91
-
-
3
Propylene
0.8395
100 000
-
-
0.86
100 000
4
Ethylene
0.0143
1137.99
0.6
1137.99
-
-
Mass
Flowrate,
(lb/hr)
Mass
Flowrate,
(lb/hr)
Mole
Fraction
Mole
Fraction
Mass
Flowrate
, (lb/hr)
42
2.5.6
Mass Balance for Separation Unit 3
8
C3H6
7
117 060.53 lb/hr
S-3
C3H8
C3H6
C3H8
9
Stream 7
Stream 9
Stream 8
No.
Component
Mole
Fraction
Mass
Flowrate,
(lb/hr)
Mole
Fraction
Mass
Flowrate,
(lb/hr)
Mole
Fraction
Mass
Flowrate,
(lb/hr)
1
Propane
0.14
17060.53
1
17060.56
-
-
2
Propylene
0.86
100000
-
-
1
100000
43
2.5.7
Smmary Table Mass Balance
Stream
Component
MassFlowrate (lb/hr)
1
Propane
104 800.56
Ethane
2 029.78
Propane
121 861.09
Ethane
2 029.78
Propane
17 060.53
Ethane
811.91
Propylene
100 000
Ethylene
1 137.99
Hydrogen
4 858.15
4
Hydrogen
4 858.15
5
Propane
17 060.53
Ethane
811.91
Propylene
100 000
Ethylene
1 137.99
Ethane
811.91
Ethylene
1 137.99
Propane
17 060.53
Propylene
100 000
8
Propylene
100 000
9
Propane
17 060.56
2
3
6
7
44
2.6
Manual Energy Balance Calculation
Unit
Operation
Substances
Moles in
(mol/h)
Moles out
(mol/h)
Temp. in
(°C)
Temp.
out (°C)
Q (kJ/h)
Globe Valve 1
Propane
Ethane
1078176.47
30669.35
1078176.47
30669.35
36
36
-3752900
Furnace
Propane
Ethane
1253683.00
30669.35
1253683.00
30669.35
36
600
115286872.6
Reactor
Propane
Propylene
Ethane
Ethylene
Hydrogen
Propane
Propylene
Ethane
Ethylene
Hydrogen
Propane
Propylene
Ethane
Ethylene
Hydrogen
Propane
Propylene
Ethane
Ethylene
Hydrogen
1253683.00
30669.35
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
175516.83
1077930.13
12267.74
18402.29
1096329.768
600
600
156385320.3
600
-103.7
162104852.8
-103.7
-129
-3612689.8
-129
-129
-
Heat
Exchanger 1
Heat
Exchanger 2
Separator 1
45
Moles in
(mol/h)
Moles out
(mol/h)
Propane
Propylene
Ethane
Ethylene
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
12267.74
18402.29
Propane
Propylene
Ethane
Ethylene
Propane
Propylene
Ethane
Ethylene
Propane
Propylene
Ethane
Ethylene
Propane
Propylene
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
12267.74
18402.29
175516.83
1077930.13
Heat
Exchanger 5
Propane
Propylene
175516.83
1077930.13
175516.83
1077930.13
Heat
Exchanger 6
Propane
Propylene
175516.83
1077930.13
175516.83
1077930.13
Unit
Operation
Pump
Heat
Exchanger 3
Heat
Exchanger 4
Separator 2
Valve
Substances
Temp. in
(°C)
Temp.
out (°C)
Q (kJ/h)
-129
-129
253588.27
-129
-10.13
17781946.57
-10.13
20
8044043.47
20
Top: -21
Bottom: 20
-23242481.41
20
20
20
38.1
25586291
38.1
40
258761.27
-93172.32
46
Unit
Operation
Separator 3
Globe Valve 2
Heat
Exchanger 7
Moles in
(mol/h)
Moles out
(mol/h)
Propane
Propylene
175516.83
1077930.13
175516.83
1077930.13
Propane
175516.83
175516.83
Substances
Propane
175516.83
Temp. in
(°C)
Temp.
out (°C)
Q (kJ/h)
40
Top: 33
Bottom: 40
26562391.2011
40
40
-409342.88
40
36
-88870.15
175516.83
47
CHAPTER 3
PROCESS SIMULATION AND HEAT INTEGRATION
3.1
Percentage Difference between Manual Calculation and HYSYS
Calculation
Table 3.1: Mass Balance
Stream No.
1
2
3
4
5
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
Mass
(lbmole/hr)
(manual)
106830.34
106830.34
123890.87
123890.87
123868.58
123868.58
4858.15
119010.43
119010.43
119010.43
1949.90
117060.53
117060.53
117060.53
100000
17060.53
17060.53
17060.53
17060.53
100000
Mass
(lbmole/hr)
(Hysys)
104100
104100
120800
120800
120800
120800
5070
115700
115700
115700
1086
114600
114600
114600
97950
16880
16880
16880
16880
97950
% Diff
2.62
2.62
2.56
2.56
2.56
2.56
4.18
2.86
2.86
2.86
79.54
2.15
2.15
2.15
2.09
1.07
1.07
1.07
1.07
2.09
48
Table 3.2: Energy Balance
Energy, kJ/h
(manual)
Energy, kJ/h
(HYSYS)
% Diff
0
0
0
Furnace
115286872.6
88440000
30.36
Reactor
158385320.3
138900000
14.03
Heat Exchanger 1
&2
165717542.6
135500000
22.30
-
-
253588.27
166300
52.49
Heat Exchanger 3
&4
25825990.04
16480000
74.91
Separator 2
2324248.41
3612500
35.66
0
0
25845052.27
15750000
64.10
26562391.2011
25110000
5.78
0
0
0
-88870.15
2489000
103.58
Equipment
Globe Valve 1
Separator 1
Pump
Valve
Heat Exchanger 5
&6
Separator 3
Globe Valve 2
Heat Exchanger 7
0
49
3.2
Heat Integration
3.2.1
Calculation of FCp
Heat exchanger 1
Stream
inlet
600°C, vap.
Stream
inlet
22.2°C,
vap.
Component
Flow rate
Ib/hr
Flow rate
mole/hr
Ethane
1428.94
21554.93
Propene
1.59
17.156589
0.099
Propane
119357.53
1227740.33
0.1181
∑FCp
147 MJ/hr. K
Cp (KJ/mol.K)
0.083
50
Heat exchanger 2
Stream inlet
Stream inlet
600°C, vap.
-129C, mixture
Flow rate
Flow rate
Ib/hr
mole/hr
Ethane
571.57
8621.972
0.071977306
Propene
97955.36
1055873.85
0.084765251
Propane
16710.05
171883.65
Ethylene
799.88
12932.96
Hydrogen
4750.25
1068790.29
∑FCp
139 MJ/hr.K
Component
Cp (KJ/mol.K)
0.099924277
0.056673009
0.029384209
51
Heat exchanger 3
Stream
inlet 20°C,
liq
Stream inlet
-128.4°C, liq
Flow rate
Flow rate
Ib/hr
mole/hr
Ethane
545.31
8225.85
0.041326623
Propene
97789.68
1054087.89
0.049096559
Propane
16687.77
171654.39
0.05463887
Ethylene
693.13
11206.91
0.033932812
Hydrogen
0.8776
197.467
0.028864532
∑FCp
61.86 MJ/hr.K
Component
Cp (KJ/mol.K)
52
Heat exchanger 4
Stream
inlet
36.08°C,
mixture
Stream
inlet 40°C,
vapor
Flow rate
Flow rate
Ib/hr
mole/hr
Ethane
152.19
2295.69
0.054581189
Propene
97789.66
1054087.72
0.066170694
Propane
16687.77
171654.39
0.076427785
Ethylene
0.7595
12.2805
0.045014187
Hydrogen
0
0
0.028847632
∑FCp
83 MJ/hr.K
Component
Cp (KJ/mol.K)
53
Heat exchanger 5
Stream
inlet
35.39°C,
vapor
Stream
inlet 10°C,
liquid
Flow rate
Flow rate
Ib/hr
mole/hr
Ethane
335.48
5060.6428
0.05248998
Propene
215540.40
2323338.49
0.06353107
Propane
22.6386
232.865907
0.073061866
Ethylene
1.6741
27.0681
0.043306918
Hydrogen
0
0
0.02884394
∑FCp
147.9 MJ/hr.K
Component
Cp (KJ/mol.K)
54
Heat exchanger 6
Stream
inlet 36°C,
vapor
Stream inlet
-42.14°C, mixture
Flow rate
Flow rate
Ib/hr
mole/hr
Ethane
0
0
0.048853226
Propene
3.5082
37.8158
0.058879756
Propane
36759.84334
378120.61
0.067127685
Ethylene
0
0
0.040291784
Hydrogen
0
0
0.028844827
∑FCp
25.384 MJ/hr.K
Component
Cp (KJ/mol.K)
55
3.2.2
Process Energy Integration
Table 3.3 below shows the data for stream that involved with heat exchanger.
Cold stream is referring to stream that needs heating while hot stream is referring to
stream that needs cooling. The ∆Tmin used in this project is 10˚C. The Tsupply and Ttarget
in table below are referring to temperature supply and temperature target.
∆Tmin = 10˚C
Table 3.3: Stream Table Data
Stream
Type
Tsupply (˚C)
Ttarget (˚C)
FCp
(MJ/hr.K)
C1
cold
22.2
600
147
C2
cold
-128.4
20
61.86
C3
C4
H1
cold
cold
hot
36.08
-42.14
600
40
36
-129
83
25.38
139
H2
hot
35.39
10
147.9
3.2.3
ΔH
(MJ/hr)
84936.60
9180.02
325.36
1983.51
82010.00
5234.18
Algorithm Table
∆Tmin chosen is divided into half which obtained 5˚C. To calculate the shifted
temperature, it is needed to adding or subtracting 5˚C for all of the temperature. For
cold stream, the actual temperature is adding by 5˚C while for hot stream, the actual
temperature is subtracting by 5˚C. Table 3.4 shows the calculated shifted temperature.
Table 3.4: Shifted Temperatures
Stream
Ts (˚C)
Tt (˚C)
C1
C2
C3
C4
H1
H2
27.2
-123.4
41.08
-37.14
595
30.39
605
25
45
41
-134
5
From table above, the pinch temperature, minimum cooling and heating
requirement are calculated and these values are calculated by using method of problem
table algorithm as shown in Figure 3.1 below.
56
T (˚C)
∆T (˚C)
∑FCpH -∑FCpC
(MJ/hr.K)
∆Hi (MJ/hr)
1st Cascade
605
6581.52
139
H1
595
10
550
-147
-8
-1470
41.08
C3
83
3.92
-91
-356.72
0.08
-8
-0.64
10.61
-33.38
-354.16
3.19
114.52
365.32
2.2
261.52
575.34
20
199.66
3993.2
41
147.9
H2
30.39
C1
147
25
5
42.14
-37.14
-123.4
C4
25.38
C2
61.86
86.26
10.6
51.76
77.14
139
-134
Figure 3.1: Algorithm Table
Hot Pinch temperature: 35.39
Cold Pinch Temperature: 25.39
-1470
5111.52
-5870
711.52
-6226.72
354.8
-6227.36
354.16
-6581.52
0
-6216.2
365.32
-5640.86
940.66
-1647.66
4933.86
533.51
7115.03
7187.61
13769.13
8661.01
15242.53
QH min
-4400
45
27.2
2nd Cascade
2181.17
6654.1
1473.4
Qc min
Pinch
57
3.2.4
Heat Exchanger Network
Then, heat exchanger network is design for the maximum energy recovery. The heat exchanger network design have above and below of
the pinch temperature. The heat exchanger network design is shows in Figure 3.2.
35.39
Above pinch
30.39
25.39
Below pinch
600
H1
H1
H3
35.39
H2
H2
-129
Add Cooler
22.2
Add Heater
H1
H2
20
-128.4
C2
H3
36.08
40
C3
Add Cooler
10
600
C1
H4
Add Heater
36
Add Heater
H4
-42.14
Figure 3.2: Heat Exchanger Network
C4
58
Table 3.5 below shows the summary of temperature for heat exchanger that obtain based
on heat exchanger network design. There are four heat exchangers that have been obtained.
Table 3.5: Summary of Temperature of Heat Exchanger
TH,in (˚C)
TH,out(˚C)
TC,in (˚C)
TC,out (˚C)
H1
600
35.39
25.39
497.3
H2
35.39
23.17
22.2
25.39
H3
35.39
-40.63
-128.4
20
H4
-40.63
-116.65
-42.14
25.39
Heat
Exchanger
After that, temperature profile of all heat exchangers is draw to determine the validity of
the heat exchange occurs. If there is any crossing in the temperature profile between the hot
and cold stream, the heat exchange is invalid. The temperature profile is shown in Figure 3.6
to 3.9.
600 ˚C
35.39˚C
497.3˚C
35.39˚C
23.17˚C
25.39˚C
25.39˚C
Figure 3.3: Temperature Profile H1
35.39˚C
20˚C
22.2˚C
Figure 3.4: Temperature Profile H2
-42.14˚C
-40.63˚C -116.65˚C
-128.4˚C
Figure 3.5: Temperature Profile H3
25.39˚C
-40.63˚C
Figure 3.6: Temperature Profile H4
59
3.2.5
Comparison of Performance Before and After Heat Integration
Table 3.6: Heating Requirement Before and After Heat Integration
Stream
Heating requirement before
Heating requirement after
heat integration (MJ/hr)
heat integration (MJ/hr)
84936.60
9180.02
325.36
1983.51
96425.49
69370.88
0
325.36
269.28
69965.52
C1
C2
C3
C4
Total
From Table 3.6,
Hot utility consumption before heat integration = 96425.49 MJ/hr
Hot utility consumption after heat integration = 69965.52 MJ/hr
Total savings =
96425.49 – 69965.52
96425.49
x 100%
= 27.4%
Table 3.7: Cooling Requirement Before and After Heat Integration
Stream
Cooling requirement before
Cooling requirement after
heat integration (MJ/hr)
heat integration (MJ/hr)
82010.00
5234.18
87244.18
10566.28
1807.25
12373.53
H1
H2
Total
From Table 3.7,
Cooling utilities consumption before heat integration = 87244.18 MJ/hr
Cooling utilities consumption after heat integration
Total savings =
87244.18 − 12373.53
87244.18
= 85.8 %
x 100%
= 12373.53 MJ/hr
60
3.2.6
Process Flow Diagram Heat Exchanger Network
C3H6
26
C1
36°C
11 bar
1
C3H8
C2H6
H2
C3H6
25 23.17°C
C2H4
15 bar
4
3
2
V-100
600 °C
1 bar
7
22.17 °C
1 bar
19.92 °C
1 bar
Mix-100
1 bar
1 bar
35.39 °C
1 bar
6
C3H8
600 °C
1 bar
C2H6
5
497.30 °C
25.39 °C HE2
HE1
Labelling
HE1
HE2
HE3
HE4
C1
C2
H1
H2
H3
Mix-100
P-100
R-100
S-100
10 °C
15 bar
H1
20
8
16
C2H4
C2H6
R-100
-26.95 °C
(To market)
20
15
HE3
24
-40.63°C
1 bar
9
T-100
C3H8
C3H6
T-100
V-101
36.08 °C
15 bar
H3
10
V-102
H2
-116.65°C
1 bar
40 °C
15 bar
43.37 °C
15 bar
21
-42.14 °C
1 bar
25.39 °C HE4
17.5 bar
H2
22
23
1 bar
19
18
17
42.79 °C
17.5 bar
36 °C
35.39 °C
15 bar
Rectifier
20 °C
17.5 bar
(To market)
T-101
C3H8
43.37 °C
15 bar
Reboiler duty
T-101
V-100
V-101
V-102
12
-129°C
1 bar
11
14
C2
-128.4°C
17.5 bar
13
-129°C
1 bar
S-100
C3H6
C2H4
C3H8
P-100
-129°C
1 bar
C2H6
Figure 3.7: Process Flow Diagram Heat Exchanger Network
Description
Heat Exchanger
Heat Exchanger
Heat Exchanger
Heat Exchanger
Cooler
Cooler
Heater
Heater
Heater
Mixer
Pump
Oleflex Reactor
Cryogenic
Separator
Distillation
Column
Splitter
Control valve
Control valve
Control valve
61
CHAPTER 4
PROCESS OPTIMIZATION
4.1
Process Optimization
Process optimization is a discipline of adjusting a process so as to optimize
some specific set of parameters without violating some constraint. The most common
goal of optimization is to minimize the cost and utilities and maximize profit.
Optimization is normally applied to improve all designs, both product and process, at
various stages in the design process. Fundamentally, there are three parameters that
can be adjusted to achieve optimal plant performance. The three parameters are
equipment optimization, operating procedures and control optimization. Firstly,
equipment optimization is to verify that the existing equipment is being used to its
fullest advantage by examining operating data to identify equipment bottleneck. Next,
optimization of operating procedures may vary widely from person to person or from
shift to shift. For example, task distribution and working procedures for processes in
the plant. Lastly, control optimization is responsible to ensure that the plant processes
are properly designed and tuned so that it can operate at optimum condition.
62
In order to optimize the process, objective function must be define. Objective
function is function that we need to minimize or maximize which might include
maximize production yield and minimize of equipment, maintenance and operating
cost. Besides, value of decision variable is also important to find as it give the optimum
value for objective function that lies within the specific constraints. Constraints is
limitation or restrictions of decision variable which can be physical, economic, policy,
or environmental constraints that can explicitly defined in terms of the decision
variable.
In this project, the objective function is to maximize the production of
propene (100000 Ib/hr) and minimize the expenses of raw material which is propane
(104800.56 Ib/hr) in which depending on the stoichiometric coefficient of
(propane/propene= 1.0478) and (hydrogen/propene= 0.048). Table 4.1 below shows
the market value of propane, hydrogen and propene.
𝐶3 𝐻8 → 𝐶3 𝐻6 + 𝐻2
Table 4.1: Market Value of Propane, Hydrogen and Propene
Chemical
Propane
Hydrogen
Propene
C3H8
H2
C3H6
1
1
2
Molar mass (Ib/Ibmol)
44.09
2.016
42.08
Mass (Ib)
44.09
2.016
42.08
Ib/Ib Propylene
1.0478
0.048
1
USD/Ib
0.195
0
0.476
RM/Ib
0.798
0
1.978
Molecular formula
Ibmol
Then, decision variable for this project are flow rate of main product
(propene), byproduct (hydrogen) and reactant (propane). There are four steps needed
in order to maximize the profit which are define all decision variables, define objective
function, define constraints and perform optimization by using suitable solution
technique (graphical linear programming software tool). For step number four, Solver
feature of Microsoft Excel is used to get the linear optimization.
63
Step 1:
Define decision variables:
P1= amount of product (Propene)
P2=amount of byproduct (Hydrogen)
R=amount of reactant (Propane)
Z=maximum profit
Step 2:
Define objective function
Maximum profit (Z) = (1.978*P1 + 0*P2)-(0.798*R)
Step 3:
Defining equality and inequality constraints:
a) Inequality constraints
i.
ii.
Propane supply
R << 104800.6 Ib/hr
Propene production P1>> 100000 Ib/hr
b) Equality constraints
i.
ii.
R= 1.0478*P1
P2=0.048*P1
c) Non-negativity constraint
R, P1, P2 ≥ 0
64
Step 4:
Optimization technique
We used solver add-in in Microsoft excel:
P1= 100019.6 Ib/hr
P2= 4800.9 Ib/hr
R= 104800.6 Ib/hr
Z= RM 197838.8/hr
As a conclusion, to maximize the profit of product, flow rate of propane is
104800.6 Ib/hr, propene is 100019.6 Ib/hr and hydrogen is 4800.9 Ib/hr. The maximum
profit of product is RM 197838.8/hr.
Figure 4.1: Solver Feature of Microsoft Excel
65
CHAPTER 5
EQUIPMENT SIZING AND COSTING
5.1
Sizing and Costing of Equipment
In this chapter, the equipment sizing is done to all equipment that is involved in
the proposed propylene production plant. Equipment sizing is a very important aspect of
process design as it enables the subsequent analysis that is involved in process design such
as mechanical design and economy analysis. The sizing involves the reactors, distillation
column, compressor, pump, and heat exchangers.
5.1.1
Oleflex Reactor
Oleflex reactor (R-100) is a type of reactor that involves a process of catalytic
dehydrogenation for the production of light olefins from their corresponding paraffin. One
of the processes of catalytic dehydrogenation is production of propylene from propane.
The Oleflex process provides high quality of propylene, which then leads to high quality
polymers. This process used Pt-Sn catalyst to promote the dehydrogenation reaction. In
this process, the Oleflex reactor operated at 600˚C and 1 bar.
66
Sizing of Reactor
Parameter
SI
Volumetric Flowrate , Q
3838.7 ft3/hr
Retention time (half-full), t
5 min
Reactor Volume, V
639.78 ft3
Vessel Inside Diameter, Di
7.41 ft
Vessel Length, L
14.83 ft
Design Type
Vertical
Material of Contruction
Low- Alloy Steel SA-387B
Type of reactor
Tubular reactor
Costing of Reactor
Cost of vessel, Cv = $ 115, 081
Cost of ladders and nozzles, CPL = $ 10, 705
Cost of purchase CP = $ 148, 802
Total cost with bare-module = 4.16 (148802) = $ 619, 017
5.1.2
Pump
The used of pump (P-100) is to provide sufficient pressure to overcome the
operating pressure of the system to move fluid like liquid at a required flow rate. In this
project, the pump was required to increase the pressure of liquid from cryogenic separator
(propane, propene, ethane and ethene) from 1 bar to 17.5 bar.
67
Sizing of Pump
Pressure inlet, P1 = 100 kPa = 14.50 psi
Pressure outlet, P2 = 1750 kPa = 253.816 psi
Pressure drop, ΔP = 1650 kPa = 239.32 psi
Volumetric flow rate, Q = 101.61 m3/hr = 447.35 gpm
Pump head, H =
𝛥𝑃 (2.31)
𝑆𝐺
=
𝛥𝑃
= 785.01 ft
ρ
Costing of Pump
Cost of pump, CP = $ 20, 995.86
Cost of motor, CP = $ 11, 880
Total cost with bare-module = (20995.86 + 11880) (3.30)
= $ 108, 492
5.1.3
Cryogenic Separator
The required temperature and pressure for the separation to occur is 129˚C and
1 bar. The used of this separator (S-100) is to remove hydrogen from the mixture vapour.
Hydrogen is removed at the distillate while other components are removed at the bottom
as the liquid.
68
Item No.
Identification
Equipment Specification Sheet
S-100
Cryogenic Separator
Process Specification / Operating Conditions
Liquid Flow Rate, FL
64893.76 Ib/hr
Vapor Flow Rate, FV
Density of liquid, ρL
55906.24 Ib/hr
43.35 Ib/ft3
Density of vapor, ρV
Surface Tension, σ
Temperature, TO
Pressure, PO
0.0116 Ib/ft3
28.6 dyne/cm
-129 °C
14.5 psig
Type
Material of Construction
Density, ρsteel
Material Factor, FM
Equipment Properties (Tower)
Vertical
Stainless Steel
0.284 lb/in.3
2.1
Maximum Allowable Stress, S
Vessel Internal Diameter, Di
Vessel Height, L
Wall thickness to withstand internal pressure, tp
Corrosion Allowance, tc
Shell Thickness, ts
Weight, W
13750 psi
12.44 ft
50 ft
0.0053 inch
0.125 inch
0.625 inch
5303.68 Ib
Cost for Tower
Cost of empty vessel, Cv
$37445.10
Added cost for Platforms and Ladders, CPL
$ 34160.95
f.o.b total purchase cost, CP
$112795.148
Bare-module cost, CBM
$469, 227.82
69
5.1.4
Distillation Column
Distillation column (T-100) is the column that separates the components of C3
and C2. Components of C2 are removed at the distillate since their boiling points are lower
than C3 components while C3 components are removed at the bottom as a liquid. This
column was operated at 20°C and 17.5 bar.
Since propane and propene have similar molecular size and physical properties,
their separation is challenging. Therefore, propane-propene splitter (T-101) is used as it
can give high purity of propene. The purified propene recovered at the top is condensed,
and the other part of it is returned as reflux while the propane is drawn off at the reboiler
and recycles to the mixer. This splitter was operated at 40°C and 15 bar.
T-100
Item No.
Identification
Equipment Specification Sheet
T-100
Distillation tower
Process Specification / Operating Conditions
Liquid Flow Rate, FL
114600 Ib/hr
Vapor Flow Rate, FV
Density of liquid, ρL
1086 Ib/hr
29.49 Ib/ft3
Density of vapor, ρV
Surface Tension, σ
Temperature, TO
1.97 Ib/ft3
7.226 dyne/cm
20 °C
Pressure, PO
253.8 psig
Type
Material of Construction
Density, ρsteel
Material Factor, FM
Equipment Properties (Tower)
Vertical
Stainless Steel
0.284 lb/in.3
2.1
Maximum Allowable Stress, S
Vessel Internal Diameter, Di
13750 psi
39.3 ft
70
Vessel Height, L
Wall thickness to withstand internal pressure, tp
Corrosion Allowance, tc
Shell Thickness, ts
66 ft
0.52 inch
0.125 inch
0.625 inch
Weight, W
22 522.22 Ib
No. of tray, NT
Material of Construction
Material Factor, FMT
Equipment Properties (Tray)
18
Stainless steel
1.58552
Cost for Tower
Cost of empty vessel, Cv
$90436.18
Added cost for Platforms and Ladders, CPL
$ 87980.27
Cost of tray, CT
$ 37035188.76
f.o.b total purchase cost
$ 3721305.21
Process Specification/ Operating Condition (Condenser)
Heat duty, Q
829800 Btu/hr
Stream
Inlet temperature, Ti
Outlet temperature, To
Pressure, P
Overall transfer coefficient,
Base cost, CB
f.o.b purchase cost, CP
Bare-module cost, CBM
Heat duty, Q
Stream
Q6
20 °C
-26.95 °C
253.8
100 Btu/(°F-ft2-hr)
Cost for Condenser
$14125
$39660
$125, 722.2
Process Specification/ Operating Condition (Reboiler)
4253000 Btu/hr
Q7
Inlet temperature, Ti
Outlet temperature, To
Pressure, P
Overall transfer coefficient,
Base cost, CB
f.o.b purchase cost, CP
Bare-module cost, CBM
20 °C
42.79 °C
253.8
75 Btu/(°F-ft2-hr)
Cost for Reboiler
$ 20454.77
$ 40909
$129, 681.53
71
Process Specification / Operating Conditions (Reflux drum)
Volumetric flowrate, Q
3612.65 ft3/hr
Retention time
0.5 hr
Volume
1806.32 ft3
Temperature, TO
Pressure, PO
20 °C
253.8 psig
Equipment Properties (Tower)
Type
Vertical
Material of Construction
Stainless Steel
Density, ρsteel
0.284 lb/in.3
Material Factor, FM
2.1
Maximum Allowable Stress, S
13750 psi
Vessel Internal Diameter, Di
8.32 ft
Vessel Height, L
33.28 ft
Wall thickness to withstand internal pressure, tp
0.11 inch
Corrosion Allowance, tc
0.125 inch
Shell Thickness, ts
Weight, W
0.5625 inch
1548.41 Ib
Cost for Reflux Drum
Cost of empty vessel, Cv
$ 811.72
Added cost for Platforms and Ladders, CPL
$2093.09
f.o.b total purchase cost, CP
$ 4444.45
Bare-module cost, CBM
$18488.91
Bare-module cost, CBM
Total Cost of Distillation Tower
$3806318.66
72
T-101
Item No.
Identification
Equipment Specification Sheet
T-101
Distillation tower
Process Specification / Operating Conditions
Liquid Flow Rate, FL
16680 Ib/hr
Vapor Flow Rate, FV
Density of liquid, ρL
97950 Ib/hr
29.82 Ib/ft3
Density of vapor, ρV
Surface Tension, σ
1.975 Ib/ft3
5.390 dyne/cm
Temperature, TO
Pressure, PO
40 °C
217.6 psig
Type
Material of Construction
Equipment Properties (Tower)
Vertical
Stainless Steel
Density, ρsteel
Material Factor, FM
Maximum Allowable Stress, S
Vessel Internal Diameter, Di
Vessel Height, L
Wall thickness to withstand internal pressure, tp
Corrosion Allowance, tc
Shell Thickness, ts
Weight, W
No. of tray, NT
Material of Construction
Material Factor, FMT
0.284 lb/in.3
2.1
13750 psi
14.2 ft
642 ft
0.163 inch
0.125 inch
0.625 inch
29533.6 Ib
Equipment Properties (Tray)
20
Stainless steel
1.58552
Cost for Tower
Cost of empty vessel, Cv
$ 107834.12
Added cost for Platforms and Ladders, CPL
$ 287443.99
Cost of tray, CT
$ 130496.2
f.o.b total purchase cost
$ 525774.31
73
Process Specification/ Operating Condition (Condenser)
Heat duty, Q
134500 Btu/hr
Stream
Q6
Inlet temperature, Ti
40 °C
Outlet temperature, To
Pressure, P
Overall transfer coefficient,
Base cost, CB
f.o.b purchase cost, CP
Bare-module cost, CBM
35.39 °C
217.6 psia
255 Btu/(°F-ft2-hr)
Cost for Condenser
$14237.98
$39898.90
$126, 479.51
Process Specification/ Operating Condition (Reboiler)
Heat duty, Q
132000 Btu/hr
Stream
Q7
Inlet temperature, Ti
40 °C
Outlet temperature, To
43.37 °C
Pressure, P
Overall transfer coefficient,
Base cost, CB
f.o.b purchase cost, CP
Bare-module cost, CBM
253.8
215 Btu/(°F-ft2-hr)
Cost for Reboiler
$ 14259
$ 28519.76
$90,407.63
Process Specification / Operating Conditions (Reflux drum)
Volumetric flowrate, Q
5915 ft3/hr
Retention time
Volume
Temperature, TO
Pressure, PO
0.5 hr
2957.5 ft3
20 °C
253.8 psig
74
Type
Material of Construction
Density, ρsteel
Equipment Properties (Tower)
Vertical
Stainless Steel
0.284 lb/in.3
Material Factor, FM
Maximum Allowable Stress, S
Vessel Internal Diameter, Di
Vessel Height, L
Wall thickness to withstand internal pressure, tp
Corrosion Allowance, tc
Shell Thickness, ts
Weight, W
2.1
13750 psi
9.8 ft
39.2 ft
0.1122 inch
0.125 inch
0.5625 inch
2006.63 Ib
Cost for Reflux Drum
Cost of empty vessel, Cv
$ 16138.02
Added cost for Platforms and Ladders, CPL
$3186.21
f.o.b total purchase cost, CP
$ 43390.10
Bare-module cost, CBM
$180, 502.82
Total Cost of Distillation Tower
f.o.b total purchase cost, CP
$ 637583.07
Bare-module cost, CBM
$2, 652 345.57
75
5.1.5
Heat Exchanger
Heat exchanger is used in both heating and cooling processes. In this process,
heat exchanger is needed to increase or decrease the temperature of the stream. There are
four heat exchangers in this production of propene.
Heat Exchanger 1 (HE1)
Sizing of Heat Exchanger 1
Hot
Cold
Tin (˚C)
600
25.39
Tout (˚C)
35.39
497.3
ΔT1 = 102.7 ˚C
𝛥𝛵𝐿𝑀
ΔT2 = 10˚C
𝑅=
600 − 35.39
497.3 − 25.39
= 1.19
𝑆=
497.3 − 25.39
600 − 25.39
10 − 102.7
=
10
𝑙𝑛( 102.7)
Based on Figure 18.15;
𝐹𝑡 = 0.55
one shell pass, one tube pass
= 0.82
𝑄 = 69370.9 𝑀𝐽/ℎ𝑟, Based on Heat Integration in Task 3
= 6.5𝑥109 𝐵𝑡𝑢/ℎ𝑟
𝑈 = 100 Btu/ 𝑜𝐹 . 𝑓𝑡 2 . hr, Based on Table 18.5, Gasoline
6.5 × 109 𝐵𝑡𝑢/ℎ𝑟
𝐴𝑖 =
100𝐵𝑡𝑢/ 𝑜𝐹 . 𝑓𝑡 2 . ℎ𝑟 × 0.55 × 103.64 𝑜𝐹
𝐴𝑖 = 11535.12𝑓𝑡 2
𝛥𝛵𝐿𝑀 = 39.80 𝑜𝐶
= 103.64 𝑜𝐹
76
1 in. OD, L
= 20 ft
= 6.096 m
1 in. triangular spacing, ID
= 0.704 in
= 0.0178 m
surface area per tube A=π*D*L
= 3.679 ft2
= 0.341 m2
𝑁=
11535.12
3.67
𝑁 = 3143.67 𝑡𝑢𝑏𝑒𝑠
From Table 18.6;
two-pass, 1.25 in trianglar pitch, shell ID=15.25 in, Do=1 in. = 0.8333ft
Surface area per pass,
𝐴𝑜 = 3.416 × 0.8333 × 20
= 8230.12𝑓𝑡 2
3143.67
2
Costing of Heat Exchanger 1
Based cost for floating head,
CB =
2
𝑒11.667−0.8709 × ln(8230.12)+0.09005 × ln(8230.12)
CB = 68503.39
𝐹𝑃 = 0.9803 + 0.018 (
𝐹𝑃 = 0.983
𝐹𝑀 = 1.75 + (
𝐹𝑀 = 2.61
14.7
14.7 2
) + 0.0017 (
)
100
100
32 0.13
)
100
FL = 1
FBM = 3.17, Based on Table 22.11
𝐶𝑝 = 1.1392 × 68503.39 × 0.983 × 2.61 × 1
𝐶𝑝 = 200394
𝐶𝐵𝑀 = 200394 × 3.17
𝐶𝐵𝑀 = $ 635 248.9
77
Heat Exchanger 2 (HE2)
Sizing of Heat Exchanger 2
Hot
Cold
Tin (˚C)
35.39
22.2
Tout (˚C)
23.17
25.39
ΔT1 = 10 ˚C
𝛥𝛵𝐿𝑀 =
ΔT2 = 1˚C
𝑅=
35.39 − 23.17
25.39 − 22.2
= 3.83
𝑆=
25.39 − 22.2
35.39 − 22.2
10 − 1
10
𝑙𝑛( 1 )
Based on Figure 18.15;
𝐹𝑡 = 0.85
one shell pass, one parallel tube
= 0.82
𝑄 = 9180.02𝑀𝐽/ℎ𝑟, Based on Heat Integration in Task 3
𝑄 = 8.7𝑥109 𝐵𝑡𝑢/ℎ𝑟
𝑈 = 100 Btu/ 𝑜𝐹 . 𝑓𝑡 2 . hr, Based on Table 18.5, Gasoline
𝐴𝑖 =
8.7 × 109 𝐵𝑡𝑢/ℎ𝑟
100𝐵𝑡𝑢/ 𝑜𝐹 . 𝑓𝑡 2 . ℎ𝑟 × 0.85 × 38.97 𝑜𝐹
𝐴𝑖 = 2626.96𝑓𝑡 2
𝛥𝛵𝐿𝑀 = 3.87 𝑜𝐶
= 38.97 𝑜𝐹
78
1 in. OD, L
= 20 ft
= 6.096 m
1 in. triangular spacing, ID
= 0.704 in
= 0.0178 m
surface area per tube A=π*D*L
= 3.679 ft2
= 0.341 m2
𝑁=
2626.96
3.67
𝑁 = 715.93 𝑡𝑢𝑏𝑒𝑠
From Table 18.6;
two-pass, 1.25 in trianglar pitch, shell ID=15.25 in, Do=1 in. = 0.8333ft
Surface area per pass,
𝐴𝑜 = 3.416 × 0.8333 × 20
𝐴𝑜 = 1874.29𝑓𝑡 2
715.93
2
Costing of Heat Exchanger 2
Based cost for floating head,
CB =
2
𝑒11.667−0.8709 × ln(1874.29)+0.09005 × ln(1874.29)
CB = 27390.39
𝐹𝑃 = 0.9803 + 0.018 (
𝐹𝑃 = 0.983
𝐹𝑀 = 1.75 + (
𝐹𝑀 = 2.61
14.7
14.7 2
) + 0.0017 (
)
100
100
32 0.13
)
100
FL = 1
FBM = 3.17, Based on Table 22.11
𝐶𝑝 = 1.1392 × 27390.39 × 0.983 × 2.61 × 1
𝐶𝑝 = 80125.52
𝐶𝐵𝑀 = 80125.52 × 3.17
𝐶𝐵𝑀 = $ 253 997.9
79
Heat Exchanger 3 (HE3)
Sizing of Heat Exchanger 3
Hot
Cold
Tin (˚C)
35.39
-128.4
Tout (˚C)
-40.63
20
ΔT1 = 15.39 ˚C
𝛥𝛵𝐿𝑀 =
ΔT2 = 87.77˚C
𝑅=
35.39 − −40.63
20 − 128.4
= 0.512
𝑆=
20 − 128.4
35.39 − −128.4
87.77 − 15.39
87.77
𝑙𝑛( 15.39)
Based on Figure 18.15;
𝐹𝑡 = 0.78
2-4 exchanger
= 0.906
𝑄 = 325.36𝑀𝐽/ℎ𝑟, Based on Heat Integration in Task 3
𝑄 = 3𝑥109 𝐵𝑡𝑢/ℎ𝑟
𝑈 = 100 Btu/ 𝑜𝐹 . 𝑓𝑡 2 . hr, Based on Table 18.5, Gasoline
𝐴𝑖 =
3 × 109 𝐵𝑡𝑢/ℎ𝑟
100𝐵𝑡𝑢/ 𝑜𝐹 . 𝑓𝑡 2 . ℎ𝑟 × 0.78 × 38.97 𝑜𝐹
𝐴𝑖 = 37𝑓𝑡 2
𝛥𝛵𝐿𝑀 = 41.58 𝑜𝐶
= 106.83 𝑜𝐹
80
1 in. OD, L
= 20 ft
= 6.096 m
1 in. triangular spacing, ID
= 0.704 in
= 0.0178 m
surface area per tube A=π*D*L
= 3.679 ft2
= 0.341 m2
𝑁=
37
3.67
𝑁 = 10 𝑡𝑢𝑏𝑒𝑠
From Table 18.6;
two-pass, 1.25 in trianglar pitch, shell ID=15.25 in, Do=1 in. = 0.8333ft
Surface area per pass,
𝐴𝑜 = 3.416 × 0.8333 × 20
𝐴𝑜 = 26.4𝑓𝑡 2
10
2
Costing of Heat Exchanger 3
Based cost for floating head,
CB =
2
𝑒11.667−0.8709 × ln(26.4)+0.09005 × ln(26.4)
CB = 17695.4
𝐹𝑃 = 0.9803 + 0.018 (
𝐹𝑃 = 0.983
𝐹𝑀 = 1.75 + (
𝐹𝑀 = 2.61
14.7
14.7 2
) + 0.0017 (
)
100
100
32 0.13
)
100
FL = 1
FBM = 3.17, Based on Table 22.11
𝐶𝑝 = 1.1392 × 17696.4 × 0.983 × 2.61 × 1
𝐶𝑝 = 51764.61
𝐶𝐵𝑀 = 51764 × 3.17
𝐶𝐵𝑀 = $ 164 093.8
81
Heat Exchanger 4 (HE4)
Sizing of Heat Exchanger 4
Hot
Cold
Tin (˚C)
-42.14
-40.63
Tout (˚C)
25.39
-116.65
ΔT1 = 74.51 ˚C
𝛥𝛵𝐿𝑀 =
ΔT2 = 66.02˚C
𝑅=
−42.14 − 25.39
−116.65 − (−40.63)
66.02 − 74.51
66.02
𝑙𝑛( 74.51 )
𝐹𝑡
1 − 50.34
√1 + 0.882 𝑙𝑛 1 − 50.34 × 0.89
𝑅 = 0.89
=
−116.65 − (−40.63)
𝑆=
−42.14 − (−40.63)
𝑆 = 50.34
(0.89 − 1) 𝑙𝑛
2 − 50.34(0.89 + 1 − √1 + 0.882 )
2 − 50.34(0.89 + 1 + √1 + 0.882 )
𝐹𝑡 = 0.79
𝑄 = 269.28𝑀𝐽/ℎ𝑟, Based on Heat Integration in Task 3
𝑄 = 2𝑥109 𝐵𝑡𝑢/ℎ𝑟
𝑈 = 100 Btu/ 𝑜𝐹 . 𝑓𝑡 2 . hr, Based on Table 18.5, Gasoline
2 × 109 𝐵𝑡𝑢/ℎ𝑟
𝐴𝑖 =
100𝐵𝑡𝑢/ 𝑜𝐹 . 𝑓𝑡 2 . ℎ𝑟 × 0.78 × 38.97 𝑜𝐹
𝐴𝑖 = 20.36𝑓𝑡 2
𝛥𝛵𝐿𝑀 = 70.18 𝑜𝐶
= 158.32 𝑜𝐹
82
1 in. OD, L
= 20 ft
= 6.096 m
1 in. triangular spacing, ID
= 0.704 in
= 0.0178 m
surface area per tube A=π*D*L
= 3.679 ft2
= 0.341 m2
𝑁=
20.36
3.67
𝑁 = 5.5 𝑡𝑢𝑏𝑒𝑠
From Table 18.6;
two-pass, 1.25 in trianglar pitch, shell ID=15.25 in, Do=1 in. = 0.8333ft
Surface area per pass,
𝐴𝑜 = 3.416 × 0.8333 × 20
𝐴𝑜 = 14.53𝑓𝑡 2
5.5
2
Costing of Heat Exchanger 4
Based cost for floating head,
CB =
2
𝑒11.667−0.8709 × ln(14.53)+0.09005 × ln(14.53)
CB = 21617.77
𝐹𝑃 = 0.9803 + 0.018 (
𝐹𝑃 = 0.983
𝐹𝑀 = 1.75 + (
𝐹𝑀 = 2.61
14.7
14.7 2
) + 0.0017 (
)
100
100
32 0.13
)
100
FL = 1
FBM = 3.17, Based on Table 22.11
𝐶𝑝 = 1.1392 × 21617.77 × 0.983 × 2.61 × 1
𝐶𝑝 = 63238.77
𝐶𝐵𝑀 = 63238.77 × 3.17
𝐶𝐵𝑀 = $ 200 466.9
83
5.1.6
Cooler
In this proposed process, there are two extra cooling utility being installed. The
cooler is operating as heat exchanger where the cooling agent is chilled brine.
Cooler 1 (C1)
Heat duty, Q = 1807.25 MJ/ hour
Hot fluid properties
Temperature inlet = 22.2 °C
Temperature outlet = 10 °C
Cooling Agent: Chilled water
Cold fluid properties
Temperature inlet = 7 °C
Temperature outlet = 32 °C
Heat transfer area = 735.5814 ft2
Length of tube = 20 ft
Number of tubes = 200.47 ≈ 201 tubes
Tube sheet layout: triangular
Tube pass: two-pass
Material (shell/tube): carbon steel/ carbon steel
Design pressure: 14.7 psig
Surface are per pass = 524.83 ft2
Based cost for floating head, CB = $ 17,065.2
F.O.B purchase cost, CP = $ 49,921.08
Cost bare-module, CBM = $ 158,249.82
84
Cooler 2 (C2)
Heat duty, Q = 10566.28 MJ/ hour
Hot fluid properties
Temperature inlet = -52.98 °C
Temperature outlet = -129 °C
Cooling Agent: Refrigerant (ethylene)
Cold fluid properties
Temperature inlet = -135 °C
Temperature outlet = -110 °C
Heat transfer area = 3345.67 ft2
Length of tube = 20 ft
Number of tubes = 911.8 ≈ 912 tubes
Tube sheet layout: triangular
Tube pass: two-pass
Material (shell/tube): carbon steel/ carbon steel
Design pressure: 14.7 psig
Surface are per pass = 2387.08 ft2
Based cost for floating head, CB = $ 30,971.67
F.O.B purchase cost, CP = $ 90,601.90
Cost bare-module, CBM = $ 287,208.01
85
5.1.7
Heater
In this proposed process, there are three extra heating utility being installed.
Heating is a process and system of raising the temperature of an enclosed space for the
primary purpose of ensuring the comfort of the occupants.
Heater 1 (H1)
Heat duty, Q = 69370.88 MJ/ hour
Hot fluid properties
Temperature inlet = 650 °C
Temperature outlet = 300 °C
Cooling Agent: Molten metals
Cold fluid properties
Temperature inlet = 128.09 °C
Temperature outlet = 600 °C
Heat transfer area = 3483.98 ft2
Length of tube = 20 ft
Number of tubes = 949.49 ≈ 950 tubes
Tube sheet layout: triangular
Tube pass: two-pass
Material (shell/tube): carbon steel/ carbon steel
Design pressure: 14.7 psig
Surface are per pass = 2485.76 ft2
Based cost for floating head, CB = $ 31,648.34
F.O.B purchase cost, CP = $ 92,581.36
Cost bare-module, CBM = $ 293,482.90
86
Heater 2 (H2)
Heat duty, Q = 325.36 MJ/ hour
Hot fluid properties
Temperature inlet = 39 °C
Temperature outlet = 33 °C
Cooling Agent: Hot water
Cold fluid properties
Temperature inlet = 36.08 °C
Temperature outlet = 40 °C
Heat transfer area = 113.60 ft2
Length of tube = 20 ft
Number of tubes = 30.96 ≈ 31 tubes
Tube sheet layout: triangular
Tube pass: two-pass
Material (shell/tube): carbon steel/ carbon steel
Design pressure: 14.7 psig
Surface are per pass = 81.05 ft2
Based cost for floating head, CB = $ 14,454.99
F.O.B purchase cost, CP = $ 42,285.39
Cost bare-module, CBM = $ 134,044.69
87
Heater 3 (H3)
Heat duty, Q = 269.28 MJ/ hour
Hot fluid properties
Temperature inlet = 38 °C
Temperature outlet = 28 °C
Cooling Agent: Hot water
Cold fluid properties
Temperature inlet = 25.39 °C
Temperature outlet = 36 °C
Heat transfer area = 83.06 ft2
Length of tube = 20 ft
Number of tubes = 22.64≈ 23tubes
Tube sheet layout: triangular
Tube pass: two-pass
Material (shell/tube): carbon steel/ carbon steel
Design pressure: 14.7 psig
Surface are per pass = 59.26 ft2
Based cost for floating head, CB = $ 14,950.04
F.O.B purchase cost, CP = $ 43,733.57
Cost bare-module, CBM = $ 138,635.43
88
CHAPTER 6
TOTAL CAPITAL INVESTMENT
6.1
Total Capital Investment
Total capital investment is total cost associated with constructing the plant. This
cost includes design, site remediation, purchasing process equipment, developing
infrastructure and contingency charges and include the raw material costs as well as labor.
There were three methods that can used in order to find total capital investment which are
Order-of-Magnitude Estimate, Study Estimate and Preliminary Estimate. In this project,
the estimation of total capital cost investment has been carried out according to the method
of Preliminary Estimate. This method based on the individual factors method Guthrie,
(1969, 1974). This method is best carried out after an optimal process
design has been developed, complete with a mass and energy balance, equipment sizing,
selection of materials of construction, and development of a process control configuration
as incorporated into a P&ID. More time is required for making a preliminary estimate than
for the preceding study estimate, but the accuracy is improved to perhaps ±20%.
89
The equation for the total capital investment by the Guthrie method is as follows:
𝐶𝑇𝐶𝐼= 𝐶𝑇𝑃𝐼 + 𝐶𝑊𝐶= 1.18 (𝐶𝑇𝐵𝑀 + 𝐶𝑠𝑖𝑡𝑒 + 𝐶𝑏𝑢𝑖𝑙𝑑𝑖𝑛𝑔𝑠 + 𝐶𝑜𝑓𝑓𝑠𝑖𝑡𝑒 𝑓𝑎𝑐𝑖𝑙𝑖𝑡𝑖𝑒𝑠) + 𝐶𝑊𝐶
The total bare-module cost, CTBM, refers to the summation of bare-module
costs for all items of process equipment, including fabricated equipment, process
machinery, spares, storage tanks, surge tanks, and computers and software. Costs for site
preparation and development, Csite, can be quite substantial for grass-roots plants, in the
range of 10–20% of the total bare-module cost of the equipment. For an addition to an
existing integrated complex, the cost may only be in the range of 4–6% of the total baremodule cost of the equipment. A detailed estimate is not normally prepared at this stage
of cost estimation.
In this Guthrie method, building costs, Cbuildings are including process buildings
and non-process buildings. A detailed estimate is also not generally made at this stage of
cost estimation. Instead, an approximate estimate is sufficient, but must consider whether
some or all the process equipment must be housed in buildings because of weather or other
conditions, and whether a grass-roots location or an addition to an integrated complex is
being considered. If the equipment is housed, the cost of process buildings may be
estimated at 10% of CTBM. If a grass-roots plant is being considered, the non-process
buildings may be estimated at 20% of CTBM. If the process is to be an addition to an
integrated complex, the non-process buildings may be estimated at 5% of C TBM. Offsite
facilities include utility plants when the company provides its own utilities, pollution
control, ponds, waste treatment, offsite tankage, and receiving and shipping facilities. This
may be added 5% of CTBM to cover other facilities. The working capital can be estimated
at 15% of the total capital investment, which is equivalent to 17.6% of the total permanent
investment.
90
There are five steps involved in Guthrie Method:
Step 1: From the process design, prepare an equipment list with equipment title, label, size,
material of construction, design temperature and pressure.
Step 2: Using the data in Step 1 with f.o.b. equipment purchase cost data, add to the equipment
list the cost, CPB and corresponding cost index, Ib of the cost data. In Guthrie method, f.o.b.
purchase cost is a base cost corresponding to a near-ambient design pressure, carbon steel as
the material of construction and a base design.
Step 3: Update the cost data to current cost index. For each piece of equipment, determine the
bare-module cost using bare-module factor, FBM. The bare module cost accounts for delivery,
insurance, taxes, and direct materials and labor for installation.
𝐶𝐵𝑀= 𝐶𝑃𝑏 (𝐼/𝐼𝑏) [𝐹𝐵𝑀 + (𝐹𝑑𝐹𝑃𝐹𝑚 −1)]
Where,
𝐹𝐵𝑀 = bare-module factor
𝐹𝑑 = equipment design factor
𝐹𝑃 = pressure factor
𝐹𝑚 = material factor
Step 4: Obtain the total bare-module cost, 𝐶𝑇𝐵𝑀, by summing the bare-module costs of the
process equipment.
Step 5: Using equation of 𝐶𝑇 (Equation 22.11, Seider 2010), estimate the total permanent
investment. Add to this an estimate of the working capital to obtain the total capital
investment.
91
6.1.1
Estimation of Total Capital Cost Investment
I = 617.62/500
Table 6.1: Summary of Bare-module Cost for All Equipment
Equipment
Item
Cp ($)
FBM
no.
Bare-module
Bare-module
Cost ($) Before
Cost ($) After
Adjusted
Adjusted
R-100
148802
4.16
619017
Pump
P-100
108, 492
3.3
108, 492
Cryogenic
S-100
112795.15
4.16
469227.82
T-100
914980.45
4.16
3806318.66
Condenser
39660
3.17
125722.2
155297.09
Reboiler
40909
3.17
129681.53
160187.81
637583.07
4.16
2652345.57
3276283.34
Condenser
39898.90
3.17
126479.51
156232.55
Reboiler
28519.76
3.17
90407.63
111675.12
HE1
200394
3.17
635248.9
HE2
80125.52
3.17
253997.9
313748.37
HE3
51764.61
3.17
164093.8
202695.23
HE4
63238.77
3.17
200466.9
247624.73
C1
49921.08
3.17
158249.82
195476.51
C2
90601.90
3.17
287208.01
354770.82
H1
92581.36
3.17
293482.90
362521.82
H2
42285.39
3.17
134044.69
165577.36
H3
43733.57
3.17
138635.43
171248.03
10393120.27
12837997.88
Oleflex
Reactor
Separator
Distillation
Column
T-101
Heat
Exchanger
Cooler
Heater
Total Bare-module Cost, CTBM ($)
764634.56
134013.66
579608.97
4701717.06
784684.85
92
Assume it is grass-roots plant; the value of CSITE is 10-20% of CTBM. Assume we take
15% of CTBM.
CSITE = 0.15 (12837997.88)
CSITE = $ 1925700
Assume it is process buildings, the value of CBUILDINGS is 10% of CTBM
CBUILDINGS = 0.10 (12837997.88)
CBUILDINGS = $ 1283800
The value of COFFSITE FACILITIES is 5% of CTBM
COFFSITE FACILITIES = 0.05 (12837997.88) + (1.5 x 107)
COFFSITE FACILITIES = $ 15641900
Use factor of 1.18 to cover a contingency and a contractor fee
CTPI = 1.18 (CTBM + CSITE + CBUILDINGS + COFFSITE FACILITIES)
CTPI = 1.18 (12837997.88 + 1925700 + 1283800 + 643399.9)
CTPI = $ 37393489
The value of CWC can be estimated 17.6% of CTPI
CWC = 0.176 (19695259)
CWC = $ 6581254
Thus,
CTCI = CTPI + CWC
CTCI = $ 37393489 + $ 6581254
CTCI = $ 43974743
93
CONCLUSION
Propylene is a major industrial chemical intermediate that serves as one of the
building blocks for an array of chemical and plastic products and also the first
petrochemical employed in the industrial scale. The main uses of refinery propylene are
in liquefied petroleum gas (LPG) for thermal use or as an octane-enhancing component in
motor gasoline. The most important derivatives of chemical and polymer grade propylene
are polypropylene, propylene oxide, isopropanol, cumene and acrylonitrile. Other
commercial derivatives include acrylic acid and esters, oxo alcohols and aldehydes,
epichlorohydrin, synthetic glycerine and ethylene-propylene copolymers. This shows that
the production of propylene has its demand in the global industry, hence a good
marketability, especially in recent years where the price of propylene in the market is
expected to continue rising as the demand increases for the chemical material. Propene
production increased in (Europe and North America only) from 2000 to 2008, it has been
increasing also in East Asia, most notably Singapore and China. Total world production
of propene is currently about half that of ethylene. About 56% of the worldwide
production of propylene is obtained as a co-product of ethylene manufacture, and about
33% is produced as a by-product of petroleum refining. About 7% of propylene produced
worldwide is on-purpose product from the dehydrogenation of propane and metathesis of
ethylene and butylenes; the remainder is from selected gas streams from coal-to-oil
processes and from deep catalytic cracking of vacuum gas oil (VGO). As Malaysia is a
part of the global market, it can be expected that prices in Malaysia to be affected by the
global prices.
94
A screening process was done based on gross profit, economic potential as well as
other factors related such as energy consumption, toxicity, safety and environmental
impacts. There are two reaction pathways suggested for the production of propylene,
which are dehydrogenation of propane, and metathesis reaction of ethylene and butene.
Based on the gross profit calculation, a dehydrogenation process would bring in a gross
profit of RM 2.72/Kg propylene with 86% conversion compared to only RM 0.70/Kg
propylene for metathesis reaction with a 90% conversion yield. It was shown that
dehydrogenation of propane reaction is a better process compared to the metathesis
reaction as it has high gross profit. Since the calculation was based on gross profit, further
analysis need to be done in order to optimize the production process of propylene via the
dehydrogenation of propane process for a sustainable plant design. After that, a process
synthesis for the production of propylene from dehydrogenation of propane was done by
following the steps that was introduced by Rudd, Powers, and Siirola. From these steps,
process flow diagram was created based on suitable operating temperature and pressure.
Then, simulation of propylene production was performed by using Aspen Hysys.
Optimization and heat integration were performed after the simulation. Sizing and costing
was also done in order to calculate the cost of all equipment. Lastly, total capital
investment is calculated. The total capital investment for this production of propylene is
$ 43974743.
95
REFERENCES
1. https://pubchem.ncbi.nlm.nih.gov/compound/Propene#section=Top
2. Aitani, A. M. (2014). Encyclopedia of Chemical Processing Propylene Production,
(January 2006). https://doi.org/10.1081/E-ECHP-120037901
3. https://patents.google.com/patent/US4753667A/en?q=propane&q=propylene&q
=splitter&oq=propane+propylene+splitter
96
APPENDICES A
Energy Balance Calculation
Reactor
>
̂𝑅
𝐻
C3H6 (g), 873K (600˚C), 1 bar
>
C3H8 (g), 873K (600˚C), 1 bar
C3H8 (g), 298K (25˚C), 1 bar
>
̂𝑟1
∆𝐻
̂𝑃
𝐻
C3H6 (g), 298K (25˚C), 1 bar
References: C3H8 (g), C3H6 (g), C2H6 (g), C2H4 (g), H2 (g) at 25°C and 1 atm
Substance
C3H8
C3H6
C2H6
C2H4
H2
nin (mol/hr)
1253683.00
30669.35
-
̂𝑅 (KJ/mol)
𝐻
̂1
𝐻
̂2
𝐻
nout (mol/hr)
175516.83
1077930.13
12267.74
18402.29
1096329.768
273.15
̂1 = ∫
𝐻
𝐶𝑝 (𝐶3 𝐻8 )
873.15
The same is done for ethane.
873.15
̂3 = ∫
𝐻
𝐶𝑝 (𝐶3 𝐻8 )
273.15
The same is done for other products.
̂𝑟1 = ∑ 𝑣𝑖 ∆𝐻
̂𝑓
∆𝐻
̂𝑓 )
̂𝑓 ) + (𝑣𝑖 )(∆𝐻
̂𝑓 )
= (𝑣𝑖 )(∆𝐻
+ (𝑣𝑖 )(∆𝐻
C3H6
H2
C2H4
̂𝑓 )
̂𝑓 )
− (𝑣𝑖 )(∆𝐻
− (𝑣𝑖 )(∆𝐻
C3H8
C2H6
̂𝑃 (KJ/mol)
𝐻
̂3
𝐻
̂4
𝐻
̂5
𝐻
̂6
𝐻
̂7
𝐻
97
̂𝑃 − 𝑛𝑖𝑛 𝐻
̂𝑅 + 𝑛∆𝐻
̂𝑟1
𝑄 = 𝑛𝑜𝑢𝑡 𝐻
𝑄 = 158385320.3 KJ/h
Pump
Component
Molar Flow Rate, F
(mol/hr)
C3H8
175516.83
C3H6
1077930.13
C2H6
12267.74
C2H4
18402.29
∑Fv = 32060954.81
Q = ∑Fv ( P)
Q = 253588.27 kJ/hr
Molar volume,v(L/mol)
Fv
21.9375
23.2486
22.1103
23.7712
3850400.458
25060366.42
271243.42
437444.516
98
APPENDICES B
99
100
101
APPENDICES C
CALCULATION OF SIZING AND COSTING
REACTOR
Q = 3838.7 ft3/hr
Retention time =5 min at half full (from Perry’s Chemical Engineering Handbook)
Volume, V = (3838.7 ft3/hr) × (
5 min×1 ℎ𝑟
60 𝑚𝑖𝑛
× 2) = 639.78ft3
Assume L/ D = 2
V = 𝜋 (D/2)2 L = (𝜋D3)/2
D = (2V/ 𝜋)1/3 = [2(639.78)/ 𝜋] 1/3 = 7.41 ft
L= 2D = 14.83 ft
Operating Pressure = 100 kPa = 14.5038 psig
Pd = exp {0.60608 + 0.91615 [ln(14.5038)] + 0.0015655 [ln(14.5038)]2} = 21.49 psig
(eqn. 22.61)
S = 13100 psi (low – alloy) (page 575)
E = 1.0
tP =
21.49 × 7.41 ×12
2 (13100)(1.0) − 1.2 (21.49)
= 0.073 in
102
Minimum wall thickness, tP = 0.375 in
tS = tP + tC = 0.073 + 0.125 = 0.198 in
W = 3.14 [7.41 + 0.198) (14.83 + 0.8 (7.41)] 0.198 (491.3) = 48263 lb
Cv = exp {7.0132 + 0.18255[ ln (48263) ] + 0.02297 [ ln (48263)]2} = $ 115, 081
CPL = 361.8 (7.41) 0.73960 (14.83) 0.70684 = $ 10, 705
Cp = FMCv + CPL = 1.2 (115081) + 10705 = $ 148, 802
Bare-Module cost
= 4.16 (148802) = $ 619, 017
PUMP
Pressure inlet, P1 = 100 kPa = 14.50 psi
Pressure outlet, P2 = 1750 kPa = 253.816 psi
Pressure drop, ΔP = 1650 kPa = 239.32 psi
Q = 101.61 m3/hr = 447.35 gpm
H=
𝛥𝑃 (2.31)
𝑆𝐺
=
𝛥𝑃
ρ
= 239.32 psi x
1𝑙𝑏/𝑖𝑛2
1 psi
x
𝑓𝑡3
43.9 lb
x
144 𝑖𝑛2
1 ft2
H = 785.01 ft
S = Q (H)0.5 = 447.35 (785.01)0.5 = 12533.89 gallon.ft0.5/min
ln S = 9.4362
CB = exp [9.7171 - 0.6019(9.4362) + 0.0519 (9.4362)2] = $ 5760.18
FT = 2.7, FM = 1.35 (Assume cast steel)
CP = FTFMCB = (2.7)(1.35)(5760.18) = $ 20995.86 for pump
𝑄 𝐻ρ
𝑔𝑎𝑙
PT = 33000 = 447.35 min x 785.01 ft x
43.9 𝑙𝑏
ft3
x
0.1334𝑓𝑡3
1 gal
x
1
33000
103
= 62.32
𝑙𝑏.𝑓𝑡
min
ln Q = 6.103
ηp = -0.316 +0.24015 (6.103) – 0.01199(6.103)2
= 0.703
PT
62.32
PB = ηp = 0.703 = 88.64
𝑙𝑏.𝑓𝑡
min
ln PB = 4.485
ηm = 0.80 + 0.0319(4.485) – 0.00182(4.485)2
= 0.906
PT
62.32
PC = ηpηm = (0.703)(0.906) = 97.85
𝑙𝑏.𝑓𝑡
min
ln Pc = 4.583
CB = exp [5.8259+0.13141(4.583)+ 0.053255 (4.583)2 + 0.028628 (4.583)3 –
0.0035549(4.583)4]
= $ 6600.31
FT = 1.8 (assume explosion-proof enclosure)
CP = FTCB = 1.8(6600.31) = $ 11, 880 for motor
FBM = 3.30
CPTotal (Pump + Motor) = (20995.86 + 11, 880) (3.30)
= $ 108, 492
104
CRYOGENIC SEPARATOR
Step 1: Extract the required data from ASPEN
Liquid Flow Rate, L
64893.76 Ib/hr
Vapor Flow Rate, G
55906.24 Ib/hr
Density of liquid, ρl
43.35 Ib/ft3
Density of vapor, ρg
Surface Tension, σ
Temperature, TO
Pressure, PO
0.0116 Ib/ft3
28.6 dyne/cm
-129 °C
14.5 psia
Step 2: Calculate the flow ratio parameter, FLG
𝐿 𝜌𝐺
𝐹𝐿𝐺 = ( ) ( )1/2
𝐺 𝜌𝐿
𝐹𝐿𝐺 = 0.0234
Step3: By taking plate spacing as 12-in, obtain the value of flooding correlation for
sieve, valve and bubble-cap trays, CSB from Figure 19.4 in textbook
𝐶𝑆𝐵 = 0.23𝑓𝑡/𝑠
Step 4: Calculate the surface tension factor
𝐹𝑆𝑇 = (
𝜎 0.20
)
20
𝐹𝑆𝑇 = 1.0741
Step 5: By taking the hole-area factor, FHA = 1, FF = 1, calculate the capacity
factor, C.
𝐶 = CSB FST FF FHA
𝐶 = 0.25
105
Step 6: Substituting the calculated value of C into Flooding Velocity equation
𝑈𝑓 = C(
𝜌𝐿 −𝜌𝐺 1/2
)
𝜌𝐺
𝑈𝑓 = 15.281 𝑓𝑡/𝑠
Step 7: Calculate the inside diameter of the distillation column
For FLG < 0.1 , the ratio (Ad/AT) can be estimated as
Ad/AT = 0.1
Assuming 80% of flooding f = 0.80, the inside diameter of distillation column can now be
calculated using the following formulation:
4𝐺
𝐷𝑇 = [
𝐴
(𝑓𝑈𝑓 )𝜋 (1 − 𝐷 ) 𝜌𝐺
𝐴𝑇
]1/2
𝐷𝑇 = 12.44 𝑓𝑡
Step 8: Extract the operating pressure and temperature from ASPEN and estimate
for the design pressure and temperature of the tower.
Po = 14.5 psia = 0 psig
For Po in the range of 0 psig to 5 psig, use a design pressure, Pd = 10 psig.
Operating temperature is -202 °F, hence the design temperature is of 50°F higher, which
is -150°F.
Step 9: Calculate the minimum wall thickness, tp
In the design temperature range of -20 to 650°F, in a corrosive environment, a commonly
used steel with a maximum allowable stress, S = 13,750 psi, since a lower temperature
have not been mentioned it is assumed that steel will be suitable. Assume the wall
thickness will be less than 1.25 inches, it then gives weld efficiency, E = 0.85.
From pressure vessel code formula,
106
𝑡𝑝 =
𝑃𝐷 𝐷𝑖
2𝑆𝐸 − 1.2𝑃𝐷
𝑡𝑝 = 0.0053 𝑖𝑛𝑐ℎ
The minimum wall thickness, tp should be 0.5 inch.
Step 10: Calculate the final shell thickness.
𝑡𝑠 = 𝑡𝑝 + 𝑡𝑐
𝑡𝑠 = 0.5 + .125 = 0.625𝑖𝑛𝑐ℎ
𝐿𝑖 = 50𝑓𝑡
Step 11: Calculate the weight of shell and the two heads
𝑊 = 𝜋(𝐷𝑖 + 𝑡𝑠 )(𝐿𝑖 + 0.8𝐷𝑖 )𝑡𝑠 𝜌
𝑊 = 5303.68 𝐼𝑏
Step 12: Calculate the purchase cost for vertical tower
𝐶𝑉 = exp{7.2756 + 0.18255[ln 𝑊] + 0.02297[ln 𝑊]2 }
𝐶𝑉 = $37445.10
Step 13: Calculate cost of platforms and ladders
𝐶𝑃𝐿 = 300.9(𝐷𝑖 )0.63316 (𝐿)0.80161
𝐶𝑃𝐿 = $ 34160.95
107
Step 14: Calculate the Total Purchase Cost
FM = 2.1 for stainless steel 316
𝐶𝑝 = 𝐹𝑀 𝐶𝑉 + 𝐶𝑃𝐿
𝐶𝑝 = $112795.148
From table 22.1, FBM = 4.16
CBM = $469, 227.82
DISTILLATION COLUMN (T-100)
Step 1-14 same with steps in cryogenic separator part
Tray stack, L = No. of trays x tray spacing
= (18-1) x 24 inches +120 +48
= 576 inches
Disengagement space bottom storage = 10 ft
Disengagement space top storage = 4 ft
𝐿𝑖 = 66𝑓𝑡
Step 15: Calculate Cost of tray
Cost of tray: CBT = 468 exp(0.1739 Di)
CBT = $ 412738.0032
Step 16: Calculate the cost of the installed trays
NT = 11 trays
FNT = 1 (NT <20)
FTT = 1.18 (Valve)
FTM = 1.401+0.0724Di (Stainless steel 316)
FTM = 4.2246
CT = NT FNT FTT FTM CBT
CT = $ 37035188.76
108
Condenser
𝑇𝐿𝑀 =
∆𝑇1 − ∆𝑇2
ln ∆𝑇1 /∆𝑇2
𝑇𝐿𝑀 = 60.2 °𝐹
𝐴=
𝑄
𝑈∆𝑇𝐿𝑀
𝐴 = 137.84 𝑓𝑡 2
𝐶𝐵 = exp{11.667 − .8709 ln 𝐴 + 0.09005 [( ln 𝐴)2 ]}
𝐶𝐵 = $ 14215
𝐹𝑀 = 1.75 + (
𝐴 0.13
)
100
𝐹𝑀 = 2.79
𝐶𝑃 = 𝐹𝑀 𝐶𝐵
𝐶𝑃 = $39660
FBM = 3.17; CBM = $125, 722.2
Reboiler
𝑇𝐿𝑀 = 60.2 °𝐹
𝐴=
𝑄
𝑈∆𝑇𝐿𝑀
𝐴 = 941.97 𝑓𝑡 2
𝐶𝐵 = exp{11.667 − .8709 ln 𝐴 + 0.09005 [( ln 𝐴)2 ]}
𝐶𝐵 = $ 20454.77
𝐹𝑀 = 2.0
𝐶𝑃 = 𝐹𝑀 𝐶𝐵
𝐶𝑃 = $40909
FBM = 3.17; CBM = $129, 681.53
109
Reflux drum
Assume L:D ratio 4:1
𝐷𝐼 = (
4𝑉
)1/3
𝜋 × 𝐿𝐷 𝑟𝑎𝑡𝑖𝑜
𝐷𝐼 = 8.32 𝑓𝑡
HEATER
Heater (H3)
a) Sizing
Hot fluid properties
Temperature inlet = 38 °C
Temperature outlet = 26 °C
Cold fluid properties
Temperature inlet = 25.39 °C
Temperature outlet = 36 °C
∆𝑇𝐿𝑀 =
=
∆𝑇2 − ∆𝑇1
∆𝑇2
)
∆𝑇1
ln(
((38 − 36) − (26 − 25.39))
(38 − 36)
𝑙𝑛 (26 − 25.39)
= 1.170°𝐶
𝑇
R= 𝑇 ℎ𝑜𝑡,𝑖𝑛
− 𝑇ℎ𝑜𝑡,𝑜𝑢𝑡
𝑐𝑜𝑙𝑑,𝑜𝑢𝑡 −𝑇𝑐𝑜𝑙𝑑,𝑖𝑛
38−26
= 36−25.39
= 1.1310
S=
𝑇𝑐𝑜𝑙𝑑,𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑,𝑖𝑛
𝑇ℎ𝑜𝑡,𝑖𝑛 −𝑇𝑐𝑜𝑙𝑑,𝑖𝑛
110
36−25.39
=38−25.39
= 0.8414
Based on graph 18.4, one shell pass, one parallel tubes, FT
= 0.9
Assume U = 100 Btu/°𝐹. 𝑓𝑡 2 . ℎ𝑟
Q = 269000.28 KJ/hour
= 254963.0384 Btu/hour
A=
𝑄
𝑈𝐹𝑇 ∆𝑇𝐿𝑀
254963.0384
= 100(0.9)(1.170)
= 83.05974415 ft2
𝐴𝑖 = 𝜋𝐷𝑖 ℎ
= 𝜋(6.096)(20)
=3.6693 ft2
Number of tubes needed, NT
NT =
83.05974415
3.6693
= 22.63628148 ≈ 23 number of tubes
Assuming tube sheet layout is two pass, shell inner diameter = 1 in.
b) Costing
𝐴𝑂 = 𝑁𝑇 𝜋𝐷𝑂 ℎ
= (23)𝜋(10.083)(20)
= 59.2617 ft2
Assuming a floating head design,
111
2)
𝐶𝐵 = 𝑒 (11.667−0.8709(ln(𝐴))+0.09005(ln(𝐴))
= 𝑒 (11.667−0.8709(ln(59.2617))+0.09005(ln(59.2617))
= $ 14, 950.04
Fp = 1 for design pressure = 14.7 psig
FL = 1 for tube length 20ft
𝐴
FM = 𝑎 + (100)𝑏
32
= 1.75 + (100)0.13
= 2.6123
Capital cost, 𝐶𝑃 = 𝐹𝑃 𝐹𝑀 𝐹𝐿 𝐶𝐵
= (1)(2.6123)(1)(14950.04)
= $ 43733.57
From Table 22.11 in textbook, FBM = 3.17
𝐶𝐵𝑀 = 𝐹𝐵𝑀 𝐶𝑃
= 3.17(43733.57)
= $ 138,635.43
The calculation for cooler is the same as heater.
2)