PROCESS SIMULATION
REFINERY PROCESSES
SAMPLER
Modelling and Optimization
John E. Edwards
Process Simulation Engineer, P & I Design Ltd
First Edition, June 2013
P&I Design Ltd
Released by
P & I Design Ltd
2 Reed Street, Thornaby TS17 7AF
www.pidesign.co.uk
Private distribution only
Copyright © P & I Design Ltd 2012
jee@pidesign.co.uk
Printed by Billingham Press Ltd, Billingham TS23 1LF
2
Process Simulation Refinery Processes
Contents
Section 1
Refinery Processes
5
Section 2
Thermodynamics
9
Section 3
Crude Column
13
Section 4
Vacuum Still
23
Section 5
Splitting and Product Purification
27
Section 6
Hydrotreater
43
Section 7
Catalytic Reformer
47
Section 8
Amine Treatment
53
Section 9
Miscellaneous Applications
57
Section 10
General Engineering Data
Section End
59
70
3
Preface
The process industry covers a broad spectrum of activities that involve the handling and treatment
of gases, liquids and solids over a wide range of physical and processing conditions. This manual
provides a comprehensive review of the fundamentals, definitions and engineering principles for the
study of processes encountered in hydrocarbon processing using steady state simulation
techniques. Applications are presented for a wide range of processing units involving design and
operations.
Process simulations are carried out using CHEMCAD™ software by Chemstations, Inc. of Houston.
This manual has been developed with the full support of Chemstations simulation engineers based
in Houston.
The simulation of crude distillation at atmospheric pressure, vacuum distillation and sour gas amine
treatment is covered in Section 13 Process Measurement and Control of the book “Chemical
Engineering in Practice” by J.E.Edwards. This manual includes these topics and extends the study
to other refinery processes including splitters, stabilizers, hydrotreaters and reformers.
Thermodynamics are reviewed with special reference to the application of pseudocomponent
curves and crude oil databases
Each topic is in the form of a condensed refresher and provides useful practical information and
data. Each section is numbered uniquely for contents and references, with the nomenclature being
section specific. The references are not a comprehensive list and apologies for unintended
omissions.
Reference is made to many classic texts, industry standards and manufacturers’ data. Information
has been mined from individual project reports and technical papers and contributions by specialists
working in the field.
.
The Author
http://uk.linkedin.com/pub/john-edwards/1b/374/924
John E.Edwards is the Process Simulation Specialist at P&I Design Ltd based in Teesside, UK.
In 1978 he formed P&I Design Ltd to provide a service to the Process and Instrumentation fields.
He has over fifty years’ experience gained whilst working in the process, instrumentation and
control system fields.
Acknowledgements
A special thanks to my colleagues at Chemstations, Houston, who have always given support in my
process simulation work and the preparation of the articles that make up this book:
N.Massey, Ming der Lu, S.Brown, D.Hill, A.Herrick, F.Justice and W.Schmidt of Germany
Also thanks to my associate P.Baines of Tekna Ltd for help with the organic chemistry topics.
4
Section 1
Refinery Processes
References
1.
2.
3.
4.
Shrieve, “Chemical Process Industries”, Chapter 37, 5th Edition, McGraw Hill, 1984.
J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process Technology”, Wiley, 2001.
G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000.
W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958.
Overview
Most refinery products are mixtures separated on the basis of boiling point ranges. The block
diagram, by API, shows overall relationship between the refining processes and refined products.
Refining is a mature, complex and highly integrated operation. Columns with a wide variety of
internals are used in many stages of the process. Fractional distillation under vacuum and pressure
conditions is used to separate components. Light ends are steam stripped and the heavy ends are
vacuum distilled at reduced the temperatures. Stabilizers are used to remove light ends, including
LPG, to reduce the vapor pressure for storage and subsequent processes. Absorbers and strippers
are used to remove unwanted components such as sulphur.
Simple distillation processes do not produce sufficient gasoline above the minimum required octane
number. This is achieved by converting heavy to light hydrocarbons using catalytic processes
including fluidic catalytic cracking (FCC), hydrotreating, hydrocracking, catalytic reforming and
alkylation.
5
Crude petroleum consists of thousands of chemical species. The main species are hydrocarbons
but there can be significant amounts of compounds containing sulphur (0-6%), oxygen (0-3.5%) and
nitrogen (0-0.6%). The main groups are:
Aliphatic or open chain hydrocarbons as detailed in the table:
Descriptor
-ane
-ene
-ol
-one
Aliphatic or open chain hydrocarbons (hc)
Properties
Class
Formula
saturated hc, unreactive paraffins
CnH2n+2
unsaturated hc, forms
olefines
CnH2n
additive compounds
acetylenes
CnH2n-2
reactive, OH replaced
alcohols, phenols
RCH3OH
additive compounds
ketones
RR1.CO
Member
C2H6
ethane
C2H4
ethylene
C2H2
acetylene
C2H5OH ethyl alcohol
(CH3)2.CO acetone
n-paraffin series or alkanes (CnH2n+2)
This series has the highest concentration of isomers in any carbon number range but only occupy
20-25% of that range and make low octane gasoline. Most straight run (distilled directly from the
crude) gasolines are predominately n-paraffins. The light ends primarily consist of propane (C3H8),
n-butane (C4H10) together with water which are defined as pure components.
iso-paraffin series or iso-alkanes (CnH2n+2)
i-butane (C4H10) is present in the light ends but these compounds are mainly formed by catalytic
reforming, alkylation or polymerization.
olefine or alkene series (CnH2n)
This series is generally absent from crudes and are formed by cracking (making smaller molecules
from larger molecules). They tend to polymerize and oxidize making them useful in forming
ethylene, propylene and butylene.
Ring compounds
Naphthene series or cycloalkanes (CnH2n)
These compounds are the second most abundant series of compounds in most crudes. The lower
members of this group are good fuels and the higher members are predominant in gas oil and
lubricating oils separated from all types of crude.
Aromatic series
Only small amounts of this series occur in most common crudes but have high antiknock value and
stability. Many aromatics are formed by refining processes including benzene, toluene, ethyl
benzene and xylene.
Lesser Components
Sulfur has several undesirable effects including its poisonous properties, objectionable odour,
corrosion, and air pollution. Sulfur compounds are removed and frequently recovered as elemental
sulfur in the Klaus process.
Nitrogen compounds cause fewer problems and are frequently ignored.
Trace metals including Fe, Mo, Na, Ni and V are strong catalyst poisons and cause problems with
the catalytic cracking and finishing processes and methods are used to eliminate them.
Salt, which is present normally as an emulsion in most crudes, is removed to prevent corrosion.
Mechanical or electrical desalting is preliminary to most crude processing.
Crude oil is classified on the basis of density as follows:
Light
less than 870 kg/m3
>31.1° API
Medium
870 to 920 kg/m3
31.1° API to 22.3° API
Heavy
920 to 1000 kg/m3
22.3° API to 10° API
Extra-heavy
greater than 1000 kg/m3
<10° API
Bitumen
Heavy or extra-heavy crude oils, as defined by the density ranges given, but with viscosities greater
than 10000 mPa.s measured at original temperature in the reservoir and atmospheric pressure, on
a gas-free basis
6
Natural Gas
Light hydrocarbon mixture that exists in the gaseous phase or in solution in crude oil in reservoirs
but are gaseous at atmospheric conditions. Natural gas may contain sulphur or other nonhydrocarbon compounds.
Natural Gas Liquids
Hydrocarbon components recovered from natural gas as liquids including ethane, propane,
butanes, pentanes plus, condensate and small quantities of non- hydrocarbons.
Atmospheric and vacuum distillations produce the different fractions as detailed in the table below.
Temperature
<30ºC
Description
Gaseous
Hydrocarbon
Density
Composition
Applications
C3H8, C4H10
Gas fuel or
enrichment
Crude Petroleum Fractional Distillation
40-70 ºC
70-120ºC
120-150 ºC
150-300 ºC
Gas oil
Naptha
Benzene
0.65
0.72
0.76
0.8
C8H18,
C9H20
C10H22,
C11H24
C12H26 to
C18H36
C5H12,
C6H14
C6H14,
C7H16
C8H18
Kerosene
General
Solvent
Solvent for
Home
solvent,
for oils,
oils, fats &
heating
aviation
fats &
varnishes
Jet fuel
spirit
varnishes
Gasoline, contains C6H14, C7H16, C8H18 40-180 ºC
>350 ºC
Heavy
oils
Residue
Asphalt
or
Bitumen
C18H38 to
C28H58
Diesel,
fuel oils
Roads,
Wax
paper
Further fractionation of the 70 to 150ºC cut is required to obtain the naptha and benzene cuts.
Vacuum distillation of the topped crude is required to obtain Light Vacuum Gas Oil (LGVO) and
Heavy Vacuum Gas Oil (HVGO)
When the difference in volatility between components is small a solvent of low volatility is added to
depress the volatility of one of the components. This process is known as extractive distillation.
Butenes are separated from butanes using this method with furfural as the extractant.
When a high volatility entrainer is used the process is known as azeotropic distillation. Anhydrous
alcohol is formed from 95% aqueous solution using benzene to free the azeotrope and high purity
toluene is separated using methyl ethyl ketone as the entrainer.
Typical Crude Oil Products Profile Ref EIA March 2004 Data
Product
Refined gallons/barrel (gal/bbl)
Gasoline
19.3
Distillate Fuel Oil (Inc. Home Heating and Diesel Fuel)
9.83
Kerosene Type Jet Fuel
4.24
Residual Fuel Oil
2.10
Petroleum Coke
2.10
Liquified Refinery Gases
1.89
Still Gas
1.81
Asphalt and Road Oil
1.13
Petrochemical Feed Supplies
0.97
Lubricants
0.46
Kerosene
0.21
Waxes
0.04
Aviation Fuel
0.04
Other Products
0.34
7
crude oil
dist
light gasses
Naptha (full range)
kerosenes
Diesels
Other high boilers
<30c
>30 <200
>150<270
>180<315
Heavy Naptha
eg Mutineer (Aus)
Paraffins
Napthenes
Aromatics
light naptha
heavy naptha
8
method 3
CH3 CH3 C6H6 Dimethyl benzene (aromatic)
3H2 Hydrogen
method 2
method 2
method 1
method 1
Fluid Catalytic Cracking FCC
and other methods
Catalytic cracking
and other methods
method 3
vol%
62
32
6
deg c
>35 <145
>140 <205
CH3 C2H5 C6H6 ethyl benzene (aromatic)
4H2 Hydrogen
C8H16 Ethyl Cyclohexane (napthene)
C6H12 Cyclohexane (napthene)
C12H26 Dodecane (paraffin)
C10H22 Decane (paraffin)
straight chain alkanes
cyclic alkanes
benzene ring structures
dist
RON 90 STANDARD
10% n heptane
C7H16
straight chain hydrocarbon
90% iso octane
C8H18
branched chain hydrocarbon (ALSO KNOWN AS 2,2,4 TRIMETHYL PENTANE)
Gasolines are compared to this mixture in relation to deflagration performance under pressure in a test engine
C8H18
CH4
C6H6 C2H5
2.5 H2
C6H6
3H2
C8H18
C4H8
C8H18
C2H4
% vol?
50
technically 2,2,4 trimethyl pentane
30
20
Dimethyl cyclohexane C8H18
Iso octane
Methane
Ethyl Benzene
Benzene
Hydrogen
iso octane
Butene
iso octane
Ethene
Gasoline
(500 components)
iso octane
Cyclopentane
Ethyl Benzene
Gasoline
Refinery Process Summary:
Section 2
Thermodynamics
References
1. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000
2. American Society for Testing and Materials (ASTM International), Standards Library
Global K and H Models
The following table gives a summary of suitable K and H models for common refinery processes.
Refinery Processing Thermodynamic Models Summary
Process
K Model
H Model (Forced)
Crude Atmospheric Distillation
Grayson Streed
Lee Kessler
Vacuum Distillation
ESSO
Lee Kessler
Hydrotreater
SRK
SRK
Sour Gas Treatment
Amine
Amine
FCC Gas Treatment
Peng Robinson
Peng Robinson
Propylene Splitter
Peng Robinson
Peng Robinson
Compression
BWRS
BWRS
Grayson Streed K model is primarily applicable to systems of non-polar hydrocarbons. It is good for
modelling hydrocarbon units, depropanizers, debutanizers, or reformer systems. The approximate
range of applicability is as follows:
Temperature Range
0 to 800°F
-18 to 430°C
Pressure Range
< 3000 psia
< 20000 kPa
ESSO K model predicts K-values for heavy hydrocarbon materials at pressures below 100 psia.
The average error for pure hydrocarbons is 8% for p* > 1 mmHg, and 30% for p* between 10E-06
and 1 mmHg according to API Technical Data Book Vol 1. It is good for modelling vacuum towers.
Lee Kessler H model is good for hydrocarbon systems.
AMINE K model is based on the Kent Eisenberg method to model the reactions with diethanolamine
(DEA), monoethanolamine (MEA), methyl diethanolamine (MDEA) being included.
The chemical reactions in the CO2-Amine system are described by the following reactions:
RR'NH2+
RR'NCOO + H2O
CO2 + H2O
HCO3H2O
↔
↔
↔
↔
↔
H+
RR'NH
HCO3CO3- H+
+ RR'NH
+ HCO3
+ H+
+ H+
+ OH-
Where R and R' represent alcohol groups. The reaction equations are solved simultaneously to
obtain the free concentration of CO2. The partial pressure of CO2 is calculated by the Henry's
constants and free concentration in the liquid phase.
The AMINE K Model in CHEMCAD treats the absorption of CO2 in aqueous MEA as a fast
chemical reaction, in other words, gas film controlled implying a very low stripping factor. However it
is known that this process is liquid film controlled since Henry’s Law controls the diffusion of CO2
into the liquid prior to chemical reaction taking place.
9
Section 3
Crude Column
Case/File Name
R3.01
Crude Column Simulations
Description
Crude Column Feed
References
1. H.Kister, “Distillation Design”, McGraw-Hill, ISBN 0-07-034909-6
2. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000
3. W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958
Process Description
The simplified process flow diagram shows the basic layout for the crude and vacuum distillation
units.
Desalted crude is preheated with the pump around and topped crude heat exchangers prior to
being heated to ~620ºF in the direct fired furnace. Above this temperature thermal decomposition
(cracking) will take place resulting in increased light ends and fouling of heat exchange surfaces
due to carbon based deposits. The following initial guidelines are suggested:
1.
2.
3.
For paraffin based crudes at moderate furnace temperatures, an estimated cracked gas
rate of 5.0 SCF/bbl (42 gal/bbl) crude oil is reasonable.
For asphalt based crude oil a cracked gas production of 2.5 SCF/bbl crude oil is suggested.
The cracked gas may be given an arbitrary composition as follows:
50 mol% methane, 40 mole% ethane, and 10 mole% propane.
The feed to the atmospheric crude tower is a mixed vapor-liquid phase of ~0.4 vapor fraction. The
vapours flow upwards and are fractionated to yield the products.
Crude towers are typically 4m diameter, 20–30m in height with 15–30 trays.
10
A typical Process Flow Diagram for a crude unit, including pump-around circuits and side strippers,
is shown. The column is modelled on the basis of theoretical stages, as opposed to actual trays, so
it is necessary to apply tray efficiency η to translate the actual trays NA to theoretical trays NT where
η=NT/NA. Note that commercial simulators provide various tray efficiency models, which are not
suitable for crude distillation columns. Tray efficiency η should be based on experience. The
relationships between NA and NT are indicated in the diagram.
The liquid product sidestreams contain light hydrocarbons which must be removed to meet the
initial boiling point specification for the products. The liquid sidestreams are fed to strippers that use
either a reboiler or steam to strip out these light components which are returned to the crude tower.
Current preference is to use reboiled side strippers for the lower boiling products to reduce the heat
load on the crude tower condenser and the sour water stripper.
Side strippers are typically 1-2m diameter, 3m in height with 4–8 trays representing 2–3 theoretical
stages. Height limitations can be met by using structured packing which has high capacity and low
HETP values as compared to trays.
Pumparound cooling circuits provide reflux to remove the latent heat from hot flash zone vapors
and condense the side products. A pump-around zone may be considered equivalent to an
equilibrium flash where equilibrium liquid is recirculated. The large flow of pump-around liquid
creates a region of constant liquid composition that eliminates fractionation. The heat removed
preheats the crude feed.
11
Section 4
Vacuum Still
Case/File Name
R4.01
Vacuum Still Simulations
Description
Vacuum Unit
Vacuum distillation is used to separate the high boiling bottoms from the crude column. The
Vacuum Unit process flow diagram is shown with distillation UnitOp 1 selected as Tower+.
The thermodynamic selection is K Model ESSO and H Model Lee Kessler.
The feed is defined by the following specification:
Feed rate
360 m3/day
Bulk gravity
0.9168 specific gravity
Feed temperature
150ºF
Feed pressure
58 psia
Distillation curve volume % based on TBP at 1 atm
12
The column specifications are:
Vacuum Column Data
Description
Specification
Number of strippers
0
Number of pumparounds
2
Number of exchangers
1
Number of side products
2
Stages
Theoretical 8 Feed 8
Column pressures
Top 30 mmHg dP 35 mmHg
Stripping Steam condition 335ºF and 115 psia
Bottom steam flow
166.67 lb mol/h
Condenser
Total
Reboiler
None
Pumparound 1
Stages
Draw-3 Return-1
Flow
276218 kg/h Phase liquid
Duty
0 MJ/h
Pumparound 2
Stages
Draw-5 Return-4
Flow
538139 kg/h Phase liquid
Duty
0 MJ/h
Side Product Draw 1
Stage
3
Flow
72 m3/h Phase liquid
Side Product Draw 2
Stage
5
Flow
213 m3/h Phase liquid
Side Heat Exchanger
Stage
8 No duty (Feed stage)
Stage Specifications
Stage
3 1 kmol/h Liquid flow
Stage
5 85 m3/h Liquid Flow
Stage
8 69 m3/h Liquid Flow
Pseudocomponent Curves allow group plots to be generated for the streams:
13
Section 5
Splitting and Product Purification
Case/File Name
R5.01
R5.02
R5.03
R5.04
R5.05
R5.06
R5.07
1.
2.
3.
4.
Splitting and Product Purification Simulations
Description
Deethanizer
Debutanizer Depropanizer
Debutanizer Reflux Depropanizer
C3 Splitter
C4 Splitter
C4 Splitter Tray Column
Kerosene Splitter
References
H.Kister, “Distillation Design”, McGraw-Hill, ISBN 0-07-034909-6
G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000
G.L.Kaes, “Practical Guide to Steady State Modelling of Petroleum Processes
H.Kazemi Esfeh and I.Aalipour mohammadi, “Simulation and Optimization of Deethanizer
Tower”, 2011 International Conference on Chemistry and Chemical Process, Singapore
Introduction
A primary activity in hydrocarbon processing involves the fractionation and purification of light ends
using columns, the most common being stabilizers, deethanizers, debutanizers and depropanizers.
A typical purification plant schematic is shown:
Deethanizer
The deethanizer removes ethane (C2H6) and lighter components which may be fed to the olefines
unit for production of ethylene (C2H4) or polyethylene or polypropylene products. Bottoms are fed to
the debutanizer. Design for C2 mole fraction or C2/C3 mole ratio in the bottoms.
Debutanizer
The debutanizer separates mixed LPG product (mostly C3’s and C4’s) and a stabilized condensate
(C5+). Design for RVP in bottoms with 12 psia being typical and reflux ratio 0.5 – 1.0
Depropanizer
The depropanizer separates propane (C3’s) as overheads from the butane (C4) to the bottoms.
Stabilizer
Stabilizers are used to remove light ends (mainly C4’s) from condensate to meet Reed Vapour
Pressure (RVP) specification for future processing or to allow storage in floating roof tanks. Design
for RVP in bottoms with 12 psia being a typical maximum value
All purification units use the bottom tray or reboiler temperature and reflux for control. The stabilizer
uses bottom tray or reboiler temperature alone as there is no condenser for reflux control. Using
these parameters in process simulation allows predicted product properties to be compared against
actual process conditions. Simulation parameters can be adjusted to match current behaviour to
provide a powerful troubleshooting tool.
14
Splitters are used extensively in hydrocarbon processing, including C2’s, C3’s, C4’s and Naphtha.
The process simulation methods used are similar to those for the purification process with the
CHEMCAD SCDS UnitOp being used.
Process
Naphtha Splitter
C2 Splitter
C3 Splitter
C4 Splitter
Tray Column Industry Practice and Efficiencies (1)
Actual Trays
Overall Efficiency
25 - 35
70 - 75
110 - 130
95 - 100
200 - 250
95 - 100
70 - 80
85 - 90
Theoretical Trays
18 - 25
105 - 125
190 - 240
60 - 68
C2 Splitter (C2H6 – C2H4)
This involves the separation of ethylene from ethane using low temperature distillation. The splitter
is normally operated at high-pressure, utilizing closed-cycle propylene refrigeration. The objective is
to obtain a high % recovery of high purity ethylene. This process is a high energy user and costly.
C3 Splitter (C3H8 – C3H6)
This involves the separation of propylene form propane. High pressure, typically 220 psia, is
needed to condense the propylene vapor at ambient temperatures around 40°C and allows the use
of cooling water on the condenser.
C4 Splitter (iC4H10 – nC4H10)
This involves the separation of i-butane form n-butane.
Naphtha Splitter
Full Range Naphthas (FRN) feed is taken from the crude unit overheads and the splitter separates
the light from the heavy naphtha. Light naphtha from the overheads is cooled against the incoming
FRN and then condensed in air fin fan coolers and used as reflux or routed to the light naphtha
stabilizer column for stabilization and recovery of light ends LPG.
The column uses a forced circuit fired reboiler system. The splitter bottoms are pumped via a heat
exchanger to recover heat from the Naphtha Hydrotreater hot reactor effluent into a fired furnace to
provide the desired reboiler duty to effect the separation of the light and heavy naphthas.
Heavy naphtha from the column bottoms is fed to the Naphtha Hydrotreater section and
subsequently the Catalytic Reformer feedstock.
The thermodynamics suitable for simulating these hydrocarbon mixtures are the equation of states
Soave – Redlich - Kwong (SRK) for pressures >1 bar and Peng Robinson for pressures >10 bar.
15
Case R5.01 Deethanizer (4)
Ethane is the primary component in the feed to olefin plants for
the production of unsaturated hydrocarbons such as ethylene.
= + Methane and ethane are to be separated from propane using
48 theoretical stages with the feed being introduced on tray15.
Integral condenser is stage 1 and reboiler stage 48.
Column top pressure is 18.33 bar but tray pressure drop was
not included.
Feed composition and conditions are shown in Stream 1.
Suitable thermodynamics are SRK or Peng Robinson.
Reference (4) indicated good agreement with both methods but
the prediction, by both methods, of ethane composition in tower
bottoms was inaccurate leading to a higher ethane recovery
than on plant
The column operating conditions are to be established to
achieve the following separation.
C2 Splitter Operating Targets
Overhead
Bottoms
Component
mole fraction
mole fraction
methane
0.241
0
ethane
0.738
0.0022
propane
0.0106
0,9914
i-butane
0
0.00555
n-butane
0
0.00085
H2S
0.000047
0
The SCDS convergence parameters were set for a distillate propane composition 1.4 reflux ratio
and a bottoms ethane composition 0.0022 mass fraction.
The column converged with a reboiler duty of 3461 MJ/h. Tray composition profile is shown.
16
Case R5.06 C4 Splitter Tray Column
The previous data has been based on an industrial fractionator; reference: Klemola and Ilme,
Ind.Eng.Chem.35, 4579 (1996) with tray specification as follows:
Column Height
Column Diameter
Number of Trays
Weir Length (side)
Weir Length (centre)
Liquid Flowpath Length
Active Area
Downcomer Area (side)
Key Tray Specifications
Downcomer Area (centre)
m
51.8
Tray Spacing
m
2.9
Hole Diameter
no
74
Total Hole Area
m
1.859
Outlet Weir Height
m
2.885
Tray Thickness
m/pass 0.967
2
Number of Valves
m
4.9
m2
0.86
Free Fractional Hole Area
m2
m
mm
m2
mm
mm
no/tray
%
0.86
0.6
39
0.922
51
2
772
18.82
SCDS simulation model is now changed to Tray Column Mass Transfer and the tray details are
entered as shown. Tray efficiency profiles were not entered but 85 to 90% is typical.
The side weir dimension is as shown in the
diagram below and is not to be confused with
side weir length.
Note that the Downcomer side area shown in
the table is for 2 passes.
Area 0.43
0.335
1.859
2 Passes
The total hole area is shown as 0.922m2 which
is in ratio to the active area of 4.9m2 giving the
free fractional hole area of 18.82%.
The simulation is now shows the following
results:
17
Section 6
Hydrotreater
Case/File Name
R6.01
Hydrotreater Simulation
Description
Hydrotreater
References
1. J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process Technology”, Wiley, 2001.
2. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000.
3. W.L.Nelson, “Petroleum Refinery Engineering”, 4th Edition, McGraw Hill, 1958.
Process (1, 2)
Hydrotreaters are used to selectively remove undesirable elements and hydrogen saturate
unsaturated components. Reactor pressures vary 500-1000psig and temperatures 550-700°F.
Hydroteating involves reaction with hydrogen to remove mainly sulfur, nitrogen and oxygen with
some hydrogenation of double bonds and aromatic rings taking place. Hydrotreating is always
applied as a pre-treatment to naphtha reforming to protect the catalyst against S-poisoning.
Hydrotreating of heavy residues is not considered here.
H2 reacts with mercaptans (1H2), thiophenes (3H2) and benzothiphenes (5H2) produce H2S
H2 reacts with pyridine (5H2) produces NH3
H2 reacts with phenols (1H2) produces H2O
Hydrotreater Hydrogen Usage and Losses (2)
Units
scf/barrel fresh feed sm3/m3 fresh
Remarks
Reactions(basis fresh feed
100 – 500
18 -89
Unsaturated components >H2
Solubility(basis fresh feed)
10 - 20
1.8 – 3.6
Purge
40 - 100
7.2 - 18
Depend H2 in makeup gas
Recycle
500 - 1500
89 - 267
Maintain H2/hydrocarbon ratio
The flow sheet is similar for all hydrotreating operations. The liquid feed stock is mixed with a
hydrogen-rich gas and preheated by exchange with the reactor effluent. The warm feed is brought
to the desired reaction temperature in a furnace and fed to the hydrotreating reactor. The reactor
effluent is cooled and the hydrogen-rich gas is separated from the liquid product. The separator
liquid is sent to a fractionator for removal of dissolved light hydrocarbon liquids and gases.
18
Case R6.01
Hydrotreater Unit
Simulation flowsheet
Simulation Parameters
Thermodynamics selection is K-SRK and L-SRK.
Pseudocomponents are created for the feed (FEEDC6+) and product (PRODC6+) streams based on
predicted molecular weight, API gravity and normal boiling point.
19
Section 7
Catalytic Reformer
Case/File Name
R7.01
Reformer Simulation
Description
Catalytic Reformer
References
1. G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1st Edition, March 2000.
2. A.Askari et al, “Simulation and Modelling of Catalytic Reforming Process”, Petroleum & Coal
ISSN 1337-7027.
Process Description
A traditional reforming process uses three fixed bed reactors in series. Endothermic
dehydrogenation reactions take place in the first two reactors requiring fired inter-heaters to raise
the temperature for the following reactor with hydrocracking reactions being significant in the final
reactor.
Reforming catalysts are subject to poisoning by hydrogen sulphide and other sulfur compounds,
nitrogen, and oxygen which are removed from the naphtha by mild hydrotreating. The primary
reformer feed stock is virgin (uncracked) naphtha from the crude distillation process and other
naphtha stocks of suitable boiling point range are acceptable after hydrotreating.
The reactions do not occur evenly through the reactors so it is the convention in simulation work to
consider all the reaction taking place in the last reactor. Reactors 1 and 2 are set up for mass
transfer with the pressure drop being entered and the isothermal mode being used to set the outlet
temperature. Using the initial reactor inlet composition for the inter-furnace duty calculations does
not result in significant inaccuracies. The final reactor is set up in adiabatic mode with the kinetic
reactions specified.
Pre-treated naphtha is combined with recycle gas with H2 composition in range 75 to 85 mole %
and preheated by exchange with the effluent from reactor 3. Typical reactor pressures and
temperature drops are shown:
Operating Data
Inlet temperature °F
Inlet pressure
psia
Measures ∆T (Typical) °F
Recycle
MMSCFD
Catalyst Volume
ft3
Reactor 1
937
413
60 (90-130)
0.5-1% Naphtha Feed
274
Reactor 2
937
394
35 (40-60)
Reactor 3
937
394
25 (1) (10-20)
411
910
Note 1 Simulation temperature drop is much larger due to the example reactions considered
The temperature drops across the reactors are monitored to track catalyst activity.
Separator parameters
The separator feed is cooled to 90 - 100°F using air and water coolers and the flash drum pressure
was run at 290 psia and with isentropic flash. The hydrogen rich gas stream is used in other refinery
operations and compressed and remainder recycled to the process where it combines with the
naphtha feed prior to the feed/effluent heat exchanger.
Operating Data
Temperature °F
Pressure
psia
H2 Purity
Separator
100
290
0.79
20
Stabilizer parameters
The liquid is feed to the stabilizer to remove the light ends. Reformer stabilizers generally have 30
to 36 actual trays with overall tray efficiencies in the range 70 to 75%. The primary function is to
strip the n-butane from the reformate product. The distillate is sent to a gas recovery plant and the
column bottoms product is stabilized reformate.
Operating Data
Number of stages
Feed tray
Feed temperature
°F
Tray 1 temperature
°F
Bottom temperature
°F
Partial condenser pressure psia
Stabilizer
36
19
297
257
446 (Simulation 488)
239
To indicate the principles of configuring the catalytic reformer the following stoichiometric equations
have been used.
Refer to Section 1 for a more detailed analysis of refinery chemistry.
It is recognised that the compositions do not represent typical conditions and is an over
simplification of the number of species and the complexity of the kinetic reactions involved.
In practices there are many reactant components and intermediate products which make it
extremely complicated to study rigorously. To reduce this complication reactants are classified into
definite groups as pseudocomponent streams.
There are many models available for the study of reaction kinetics including Langmuir-Hinshelwood,
Arrhenius and Smith. The Smith model considers the following groupings:
Naphthene + H2
→
Paraffin
Naphthenes
→
Aromatics + 3H2
Hydrocracking of paraffin
Hydrocracking of naphthenes
Pseudocomponent groupings will include specific boiling ranges such as for C6 to C11 paraffins, C6
to C11 naphthenes, benzene, toluene and C8 to C11 aromatics.
21
Case 7.01 Catalytic Reformer
Flowsheet is shown:
Thermodynamics selected:
K-Peng Robinson
The configuration for catalytic Reactor 3 is shown.
22
H-Peng Robinson
Section 8
Amine Treatment
Case/File Name
R8.01
Vacuum Still Simulations
Description
Sour Gas Treatment
References
1. A.Kohl and R.Nielsen, “Gas Purification”, Gulf Publishing , 5th Edition, 1997
Process
Chemical absorption of CO2 and H2S with amines provides the most cost effective means of
obtaining high purity vapor from sour gases in a single step. The process is well established for
refinery gas sweetening which are carried out at high pressures. Several alkanolamines such as
MEA (monoethanolamine), DEA (diethanolamine) and MDEA (methylydiethanolamine) have been
used, with the selection being determined by the application.
The speciality amine aqueous solution strength can vary in the range 15 to 50% and can have a
significant effect on the process economics. Generic amines, such as MEA and DEA, are more
corrosive and strengths are limited to 30%. The higher the solution strength the liquid circulation
requirement is reduced. Steam consumption is highly dependent on this selection, with lower
concentrations requiring more steam. The boiling point for regeneration increases at higher MEA
concentrations which greatly increases the rate of corrosion of common metals. Also MEA tends to
degrade as the temperature rises, increasing the replacement expense and involving the removal of
degradation products.
MEA has a substantially higher vapor pressure than other amines and a water wash at the top of
the absorber can be used to minimize amine losses. The Lean Amine Feed temperature can have a
significant influence on the amine losses.
A typical gas sweetening flowsheet using amine solution is shown. It consists of an absorber in
which cooled lean solvent flows downward contacting with the upward flowing gas to be treated.
The rich liquid leaves the absorber at a higher temperature due to the heats of solution and reaction
and is preheated with stripper bottoms prior to being fed to the reboiled stripping column. The
overhead stripped gas is cooled to remove water vapor which is returned to the column to maintain
the water balance. If the gas stream to be treated contains condensable hydrocarbons the lean
amine temperature should be above the dew point temperature to prevent condensation of an
immiscible hydrocarbon liquid which will promote foaming in the absorber.
23
Section 9
Miscellaneous Application
Case/File Name
R9.01
Miscellaneous Application Simulations
Description
Biodiesel Blending
Biodiesel Blending
Diesel is blended with Methyl Ester (ME) during ship offloading. The diesel composition in the ester
can vary from 0 to 20%v/v and the product blend ester composition is in the range 5% to 15%v/v.
The ship discharge flow can vary from 800m3/h to 1000 m3/h at a maximum pressure of 10 barg.
The mass balance on the streams give:
DE
E
DE + E
VE
DS
DS + DE + E
VP
Where:
DE
E
VE
DS
VP
Diesel in Methyl Ester Flow
(m3/h)
Methyl Ester Flow
(m3/h)
Methyl Ester Volume Fraction
Diesel Flow from Ship
(m3/h)
Bio-diesel Product Volume Fraction
E
E
and
VP =
E + DE
E + DE + DS
E 1 − VE
and substituting for DE leads to the following:
DE =
VE
DE + E
1
V P DS
and
=
E =
(V E V P − 1)
1 − VP VE
DS
VE =
We have:
(
Rearranging gives:
(
Methyl Ester
%
100
80
)
)
Ester Blend to Ship Flow Ratios
Product Blend
VE / VP
%
15
10
5
15
10
5
6.667
10.0
20.0
5.33
8.0
16.0
Ship to Ester
Blend
Flow Ratio
0.176
0.111
0.053
0.231
0.143
0.067
A process control system would set the blender flow ratios by entering the ME blend (VE) and Final
Product blend (VP). The ME blend flow (E+DE) required for a “wild” Ship Discharge flow (DS) is
calculated. The flow ratio controller would manipulate the ME flow to achieve the desired ratio.
In the simulation the blend actual ME component standard liquid volume fraction and the product
ME standard liquid volume fraction are set in the appropriate controllers.
24
Case R9.01 Biodiesel Blender Simulation
The physical property data used are shown. The ship diesel temperature can vary in the range
15ºC to 40ºC. Thermodynamics used were K- UNIFAC and L- Latent Heat.
Fluid
Diesel
Methyl Ester !00%
Methyl Ester 80%
Fluid Physical Property Data at 15ºC
Density
kg/m3
807.15
876.43
864.74
Viscosity
cps
2.78
7.48
5.72
The controllers are configured as shown:
The simulation results are in accordance with the theory developed above.
The maximum and minimum ester blend flowrates obtained are 230.7 and 42.1 m3/h.
25
Section 10
General Engineering Data
Contents
Units
Refinery Process Overview
Commercial Steel Pipe ANSI B36.10:1970 & BS 1600 Part 2: 1970
Typical Overall Heat Transfer Coefficients
Typical Fouling Resistance Coefficients
Heuristics for Process Equipment Design
Process Simulation Procedures and Convergence
References
1. Crane Co., “Flow of Fluids Through Valves, Fittings and Pipes”, Publication 410, 1988
Units
Volume
The basic measurement for crude oil liquid volume is referred to as a barrel (bbl). CHEMCAD unit
converter feature by selecting “Fn f6” allows conversion between units as shown:
Gas Constant R
8.314 J/K-mol
1.986 Btu/R-lbmol
0.73 ft3 atm/ R lb mol
API gravity formulae
API gravity (API ρ) of petroleum liquids is determined from specific gravity (SG) at 60°F:
=
141.5
131.5
The specific gravity of petroleum liquids can be derived from the API gravity:
=
141.5
+ 131.5
Heavy oil with a specific gravity of 1.0 (density water at 60°F) has an API gravity of:
141.5 131.5 = 10°
Crude oil is often measured in metric tons (1000kg). The number of barrels per metric ton for a
given crude oil based on its API gravity is calculated from:
= 1
141.5
0.159!
+ 131.5
Where 1 bbl = 0.159 m3
A metric ton of West Texas Intermediate 39.6° API would contain about 7.6 barrels.
26
Reid Vapor Pressure (RVP)
A measure of gasoline volatility being defined as the absolute vapor pressure exerted by a liquid at
100°F(37.8 °C) as determined by the test method ASTM-D-323. The test method applies to volatile
crude oil and volatile non-viscous petroleum liquids.
Cetane Index
Based on the density and distillation range ASTM D86 of a hydrocarbon using two methods ASTM
D976 and D4737 (ISO 4264). Cetane index in some crude oil assays is often referred to as Cetane
calcule, while the cetane number is referred to as Cetane measure.
Aniline Point
Defined as the minimum temperature at which equal volumes of aniline and oil are miscible to give
an estimate of the content of aromatic compounds in the oil. The lower the aniline point, the greater
is the content of aromatic compounds.
VABP and MeABP
Petroleum fractions are cuts with specific boiling point ranges, API gravity and viscosity. Each cut
can be divided into narrow boiling fractions called pseudo-components where the average boiling
point can be estimated as either mid-boiling point or mid-percentage boiling point. The TBP curve is
divided into an arbitrary number of pseudo-components or narrow boiling cuts. Since the boiling
range is small both mid-points are close to each other and can be considered as the VABP or
MeABP for that pseudo-component.
Five different average boiling points can be estimated on the distillation curve. The volume average
boiling point (VABP) and the mean average boiling points (MeABP) are the most widely used.
VABP is calculated from the ASTM D86 distillation and is the average of the five boiling point
temperatures (°F) at 10, 30, 50, 70 and 90% distilled:
"# =
MeABP is calculated from:
$%& + $'& + $(& + $)& + $*&
5
+# = "# ∆
Where ∆ is given by:
∆ = 0.94402 0.008650"# 321&.) + 2.997913&.'''
3 =
$*& $%&
90 10
MeABP (°R) is used in the definition of the Watson K which is given by:
4=
+#
Factors Note 1
10-12
10-9
10-6
10-3
10-2
10-1
101
102
103
106
109
1012
%/'
Prefix
pico
nano
micro
milli
centi
deci
deca
hecto
kilo
mega
giga
tera
E-12
E-09
E-06
E-03
E-02
E-01
E01
E02
E03
E06
E09
E12
Symbol
p
n
μ
m
c
d
da
h
k
M Note 2
G
T
Note 1 Tip for setting power, make equal to number 0’s so 0.00001 = 10-5 and 100000 = 105
Note 2 Refinery industry practice sometimes uses MM to signify 106
27
Heuristics for Process Equipment Design
In modelling, “Rules of Thumb” or heuristics based on experience, are used for estimating many
parameters before more specific data is available.
Piping Design
Industry practice for initial design of piping systems is based on economic velocity or allowable
pressure drop ∆P/100ft. Once detailed isometrics are available the design will be adjusted to satisfy
local site conditions.
Reasonable Velocities for Flow of Fluids through Pipes (Reference Crane 410M)
Reasonable Velocities
Pressure Drop
Service Conditions
Fluid
m/s
ft/s
kPa / m
Boiler Feed
Water
2.4 to 4.6
8 to 15
Pump Suction & Drain
Water
1.2 to 2.1
4 to 7
Liquids pumped
General Service
1.0 to 3.0
3.2 to 10
0.05
Non viscous
Saturated Steam
Heating Short Lines
20 to 30
65 to 100
0 to 1.7 bar
Saturated Steam
Process piping
30 to 60
100 to 200
>1.7 bar
Superheated Steam
Boiler and turbine leads
30 to 100 100 to 325
14 and up
Process piping
Gases and Vapours
15 to 30
50 to 100 0.02%line pressure
Process piping
Liquids gravity flow
0.05
Reasonable velocities based on pipe diameter
Process Plant Design, Backhurst Harker p235
Pump suction line for
Pump discharge line for
Steam or gas
d in
d in
d in
(d/6 + 1.3) ft/s
(d/3 + 5) ft/s
20d
ft/s
d mm
d mm
d mm
(d/500 + 0.4) m/s
(d/250 + 1.5) m/s
0.24d m/s
Heuristics for process design
Reference W.D.Seader, J.D.Seider and D.R.Lewin, “Process Design Principles” are also given:
Liquid Pump suction
Liquid Pump discharge
Steam or gas
(1.3 + d/6) ft/s
(5.0 + d/3) ft/s
(20d)
ft/s
0.4 psi /100 ft
2.0 psi /100 ft
0.5 psi /100 ft
Air for combustion, unless otherwise stated, is at ISO conditions of 15°C, 1.013 bar and 60%
relative humidity.
Air for compression is defined at Free Air Delivery(FAD) conditions of 20°C, 1 bar and dry.
Pumps
Centrifugal pumps: single stage for 15-5000 gpm, 500 ft max head.
Centrifugal pumps: multistage for 20 – 11,000 gpm, 5500 ft max head.
Efficiency 45% at 100 gpm, 70% at 100 gpm and 80% at 10,000 gpm.
28
Process Simulation Procedures and Convergence
Steady state simulation proves the capability to achieve stable and reproducible operating
conditions with acceptable product purity, yield and cycle times to satisfy the commercial
requirements and the safety and environmental issues for the regulatory authorities.
A process simulation involves taking the input stream flow rates, compositions and thermodynamic
conditions, performing a series of iterative calculations as the streams are processed through Unit
Operations and recycles, finally leading to the output stream flow rates, compositions and
thermodynamic conditions.
Products
Feeds
Recycles
The chart below shows the basic steps involved in setting up a steady simulation.
It is recommended that the SCDS UnitOp is used for building fully integrated models because it has
a greater number of connection points.
SCDS 1 and SCDS 21-24 icons are the most developed having built in dynamic vessels and control
loops. However for our initial exercise we will use SCDS Column 1 icon.
For refinery operations the Tower UnitOp is more suitable as it includes pump around and stripping
facilities.
The stage numbering convention in CHEMCAD is from top to bottom, 1 to N. A stage is considered
the space above a plate. If a condenser is present it is stage 1; if a reboiler is present it is stage N.
To model a column which has ten stages plus condenser and reboiler 12 stages
(10+condenser+reboiler = 12) must be specified.
If a condenser is present, the feed must not enter stage 1, as that is the reboiler. Top stage feeds
should enter stage 2, the top stage (plate), if a condenser is present. Likewise, if a reboiler is
present a bottom plate feed is connected to stage (N-1), not stage N.
Typically the user has a product specification, mass fraction of a key component in either the
bottoms or tops, for a column design or to achieve with an existing column.
Converging a column model in simulation is similar to converging a column in the real world; it is
difficult to go directly to high purity separation. It is best to start with an easy target, such as reflux
ratio and bottoms flowrate. Once the column is converged to this simple specification, we ‘tighten’
the specifications toward the target specification. Use the following procedure:
29
1.
2.
3.
4.
5.
6.
Set up the column: number of stages, condenser, reboiler, operating pressure.
Generate TPxy and RCM plots to verify that the target is thermodynamically feasible with
the selected VLE K model.
On the SPECIFICATIONS page, set ‘loose’ specifications such as ‘Reflux Ratio’ and ‘
Bottoms Flowrate’ or Reboiler Heat Input.
Run the column and converge. Change the specifications if necessary.
Go to the CONVERGENCE page of the column dialog. Set the initial flag to 0 Reload
Column Profile. This setting instructs CHEMCAD to use the current converged profile as its
starting point (initial conditions) in iterative calculations.
On the SPECIFICATIONS page change to more tight specifications. Run the column.
If the column converges, tighten the specifications and run again. If the column fails to converge, do
not save the profile of the failed attempt. Relax the specifications and run the column again.
Repeat from step 5 until you reach the target.
Often, it is difficult to obtain the first convergence on a column. If the column is run with no
condenser or reboiler, one does not have the option of ‘loose’ specifications. If the column has a
condenser or reboiler, relaxing specifications does not always help.
1.
2.
3.
4.
5.
On the convergence page of the column dialog, specify estimates if you can make
reasonable estimates. Note that a bad guess will make the column more difficult to
converge than no estimate.
Remove non key components from the feed(s) to obtain the first convergence. Now set the
initial flag to 0 Reload Column Profile, return the other components, and run the unit again.
Specify a larger number of iterations on the convergence page of the column dialog. The
default is 50, but possibly 52 iterations will find the answer.
Try an alternate column model. If you are currently using the SCDS try the same separation
with a TOWER or vice versa. The two models use different mathematical models; often one
will find an answer in 10 iterations while the other is difficult to converge. It is not possible to
obtain different answers with the columns; the models are numerical methods to find a
stable composition profile.
Consider a partial condenser. If you have a condenser present but have a significant
amount of light ends, you may have difficulty converging the column. The default condenser
type, total, requires that no vapor leaves stage 1. If light ends are present, this may not be
possible without cryogenic temperatures. Changing condenser mode to partial allows the
light ends gases to slip past the condenser.
30