PROCESS SIMULATION REFINERY PROCESSES SAMPLER

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PROCESS SIMULATION


REFINERY PROCESSES


SAMPLER



Modelling and Optimization






John E. Edwards


Process S
imulation Engineer, P & I Design Ltd







First Edition, June 2013

P&I Design Ltd

2


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






3


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


Ge
neral Engineering D
ata





59




Section End







70







4


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
provid
es a comprehensive review of the
fund
amentals, definitions and engineering

principle
s 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 o
f 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 ext
ends 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 unintend
ed 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 Au
thor

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 De
sign 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 assoc
iate P.Baines of Tekna Ltd for help with the organic chemistry topics.



5


Section 1

Refinery Processes


References


1.

Shrieve, “Chemical Process Industries”, Chapter 37, 5
th

Edition, McGraw Hill, 1984.

2.

J.A.Moulijn, M.Makkee, A.Van Diepen, “Chemical Process
Technology”, Wiley, 2001.

3.

G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1
st

Edition, March 2000.

4.

W.L.Nelson, “Petroleum Refinery Engineering”, 4
th

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 use
d 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 minim
um 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.


6


Crude petroleum consists of thous
ands 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 ta
ble:


Aliphatic or open chain
hydrocarbons (hc)

Descriptor

Properties

Class

Formula

Member

-
ane

saturated hc, unreactive

paraffins

C
n
H
2n+2

C
2
H
6

ethane

-
ene

unsaturated hc, forms
additive compounds

olefines

C
n
H
2n

C
2
H
4

ethylene

acetylenes

C
n
H
2n
-
2

C
2
H
2

acetylene

-
ol

reactive, OH replaced

alcohols, phenols

RCH
3
OH

C
2
H
5
OH ethyl alcohol

-
one

additive compounds

ketones

RR
1
.CO

(CH
3
)
2
.CO acetone


n
-
paraffin series or alkanes (C
n
H
2n+2
)


This series
has

the highest concentration of isomers in a
ny carbon number range but
onl
y 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 cons
ist of propane

(C
3
H
8
), n
-
butane
(C
4
H
10
) together with water which are defined as pure components.


iso
-
paraffin series or iso
-
alkanes (C
n
H
2n+2
)

i
-
butane (C
4
H
10
) is present in the light ends but these compounds are mainly formed by catalytic
reforming, alk
ylation or polymerization.


o
lefin
e

or alkene series (C
n
H
2n
)

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 (C
n
H
2n
)

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 predomina
nt 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 ben
zene, 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
s
ulfur 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 follo
ws:

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 viscositi
es greater
than 10000 mPa.
s measured at original temperature in the reservoir and atmospheric pressure
, on a
gas
-
free basis


7


Natural Gas


Light hydrocarbon mixture

that exist
s
in the gaseous phase or i
n solution in crude oil in reservoirs but
are gaseous at atmospheric conditions. Natural gas may contain sulphur or other non
-
hydrocarbon
compounds.



Natural Gas Liquids


H
yd
rocarbon components
recovered

from natural gas as liquids including
ethane, propa
ne, butanes,
pentanes plus, condensate and small quantities of non
-

hydrocarbons.



Atmospheric and vacuum distillations produce the different fractions as detailed in the table below.



Crude Petroleum Fractional Distillation

Temperature

<30ºC

40
-
70 ºC

70
-
120ºC


120
-
150 ºC

150
-
300 ºC

>350 ºC

Residue

Description

Gaseous
Hydrocarbon

Gas oil

Naptha

Benzene

Kerosene

Heavy
oils

Asphalt
or
Bitumen

Density


0.65

0.72

0.76

0.8



Composition

C
3
H
8
, C
4
H
10

C
5
H
12
,
C
6
H
14

C
6
H
14
,
C
7
H
16

C
8
H
18

C
8
H
18
,
C
9
H
20

C
10
H
22
,
C
11
H
24

C
12
H
26

to
C
18
H
36

C
18
H
38
to
C
28
H
58


Applications

Gas fuel or
enrichment

General
solvent,
aviation
spirit

Solvent
for oils,
fats &
varnishes

Solvent for
oils, fats &
varnishes

Home
heating

Jet fuel

Diesel,
fuel oils

Roads,
Wax
paper


Gasoline,
contains C
6
H
14
, C
7
H
16
, C
8
H
18

40
-
180 ºC





Further
fractionation of the 70 to 150
ºC 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 separate
d using methyl ethyl ketone as the entrainer.


Typical Crude Oil Products Profile Ref EIA March 2004 Data

Product

Refined gallons/b
arrel

(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


8



Refinery
Process S
ummary
:


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
crude oil
dist
light gasses
<30c
deg c
Naptha (full range)
>30 <200
dist
light naptha
>35 <145
kerosenes
>150<270
heavy naptha
>140 <205
Catalytic cracking
Gasoline
Diesels
>180<315
and other methods
Other high boilers
Heavy Naptha
Fluid Catalytic Cracking FCC
Gasoline
eg Mutineer (Aus)
vol%
and other methods
(500 components)
% vol?
straight chain alkanes
Paraffins
62
iso octane
50
technically 2,2,4 trimethyl pentane
cyclic alkanes
Napthenes
32
Cyclopentane
30
benzene ring structures
Aromatics
6
Ethyl Benzene
20
C10H22
Decane (paraffin)
iso octane
C8H18
method 1
Ethene
C2H4
C12H26
Dodecane (paraffin)
iso octane
C8H18
method 1
Butene
C4H8
C6H12
Cyclohexane (napthene)
Benzene
C6H6
method 2
Hydrogen
3H2
C8H16
Ethyl Cyclohexane (napthene)
Ethyl Benzene
C6H6 C2H5
method 2
2.5 H2
CH3 C2H5 C6H6
ethyl benzene (aromatic)
Iso octane
C8H18
4H2
Hydrogen
method 3
Methane
CH4
CH3 CH3 C6H6
Dimethyl benzene (aromatic)
Dimethyl cyclohexane
C8H18
3H2
Hydrogen
method 3

9


Section 2

Thermodynamics


References


1.

G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1
st

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 g
ood for
mode
l
ling hydrocarbon units, depropanizers, debutanizers, or reformer systems.

The approximate
range of
applicability is as follows:




Temperature Range















Press
ure Range



0 to 800°
F
























<

3000 psia



-
18 to 430°C
























< 20000 kPa


ESSO K model predicts K
-
values f
or heavy hydrocarbon materials
at pressur
es below 100 psia.
The
a
verage 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 g
ood for mode
l
ling vacuum towers.


Lee
Kes
s
ler H model is good for hydrocarbon sys
tems
.


AMINE K model is based on

the Kent Eisenberg method to model the reactions

with d
iethanolamine
(DEA),
m
onoethanolamine (MEA),
m
ethyl diethanolamine (MDEA)

being included.



The chemical reactions in the CO
2
-
Amine system are described by the followin
g reactions:


RR'NH2
+




H
+


+ RR'NH


RR'NCOO + H2O



RR'NH


+ HCO3

CO2 + H2O




HCO3
-


+ H
+

HCO3
-





CO3
-

-


+ H
+


H2O





H
+


+ OH
-



Where
R and R' represent alcohol groups. The reaction equations are
solved simultaneously to ob
tain
the free concentration of
CO
2
. The partial pressure of CO
2

is calculated by the Henry's constants and
free concentration in the liquid phase.


The AMINE K Model in CHEMCAD treats the absorption of CO
2

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 CO
2

into the liquid
prior to chemical reaction taking place.






10


Section 3

Crude Column


Crude Column Simulations

Case/File Name

Description

R3.01

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, 1
st

Edition, March 2000

3.

W.L.Nelson, “Petroleum Refinery Engineering”, 4
th

Edition, McGraw Hill, 1958


Process Description


The simplified process flow diagram shows the basic layout for the crude and v
acuum 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 ex
change surfaces due to carbon
based deposits.
The following initial guidelines are suggested:


1.

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.


2.

For asphalt

based crude oil a cracked gas production of 2.5 SCF/bbl crude oil is suggested.

3.

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 mi
xed 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.


11


A typical Process Flow Diagram for a crude unit, including pum
p
-
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 N
A

to theoretical trays N
T

where
η
=
N
T
/
N
A
. Note that commercial simulators
provide various t
ray efficiency models,
which are
not suitable

for crude distillation columns
. T
ray efficiency
η

should be based on experience. The relationships
between N
A

and N
T

are indicated in the diagram.




The liqu
id 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 wh
ich 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 h
eight 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.


Pump
around 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 l
arge flow of pump
-
around liquid creates a region of
constant liquid composition
that eliminates fractionation. The heat removed preheats the crude feed.




12



Section 4

Vacuum Still


Vacuum Still Simulations

Case/File Name

Description

R4.01

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 m
3
/day

Bulk gravity


0.9168 specific gravity

Feed temperature

150ºF

Feed pressure


58 psia


Distillation curve volume % based on TBP at 1 atm






















13



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 m
3
/h Phase liquid

Side Product Draw 2

Stage

5

Flow

213 m
3
/h Phase liquid

Side Heat Exchanger

Stage

8 No duty (Feed stage)

Stage Specifications

Stage

3 1 kmol/h Liquid flow

Stage

5 85 m
3
/h
Liquid Flow

Stage

8 69 m
3
/h Liquid Flow



Pseudocomponent Curves allow group plots to be generated for the streams:




14



Section 5

Splitting and
Product Purification


Splitting and Product Purification Simulations

Case/File Name

Description

R5
.01

Deethanizer

R5.02

Debutanizer Depropanizer

R5.03

Debutanizer Reflux Depropanizer

R5.04

C3 Splitter

R5.05

C4 Splitter

R5.06

C4 Splitter Tray Column

R5.07

Kerosene Splitter


References

1.

H.Kister, “Distillation Design”,
McGraw
-
Hill, ISBN 0
-
07
-
034909
-
6

2.

G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1
st

Edition, March 2000

3.

G.L.Kaes,

Practical Guide to Steady State Modelling of Petroleum Processes

4.

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, deeth
anizers, debutanizers and depropanizers. A
typical purification plant schematic is shown:


Deethanizer

The deethanizer

removes ethane (C
2
H
6
) and lighter components which may be fed to the olefines unit
for production of ethylene (C
2
H
4
) or polyethylene or
polypropylene products. Bottoms are fed to the
debutanizer.

Design for C
2

mole fraction or C
2
/C
3

mole ratio in the bottoms.


Debutanizer

The debutanizer separates mixed LPG product (mostly C
3
’s and C
4
’s) and a stabilized condensate
(C
5
+). Design for RVP in

bottoms with 12 psia being typical and reflux ratio 0.5


1.0


Depropanizer

The depropanizer separates propane (C
3
’s) as overheads from the butane (C
4
) to the bottoms.


Stabilizer

Stabilizers are used to remove light ends (mainly C
4
’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 re
flux 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

15


process conditions.
Simulation parameters can be adjusted to match current behaviour to provide a
powerful troubleshooting tool.


16


Splitters are used extensively in hydrocarbon processing, including C
2
’s, C
3
’s, C
4
’s

and Naphtha. The
process simulation methods used are similar to those for the purification process with the CHEMCAD
SCDS UnitOp being used.


Tray Column Industry Practice and Effici
e
n
cie
s
(1)

Process

Actual Trays

Overall Efficiency

Theoretical Trays

Naphtha Splitter

25
-

35

70
-

75

18
-

25

C2 Splitter

110
-

130

95
-

100

105
-

125

C3 Splitter

200
-

250

95
-

100

190
-

240

C4 Splitter

70
-

80

85
-

90

60
-

68


C
2

Splitter

(C
2
H
6



C
2
H
4
)


This involves the separation of ethylene from ethane

using
low temperature
distillation. The splitter is
normally

operated at high
-
pressure,

utilizing closed
-
cycle propylene refrigerati
on. The
objecti
ve is

to
obtain a
high % recovery of high purity ethylene
.
This process is a high energy user and costly.


C
3

Splitter (C
3
H
8


C
3
H
6
)


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.


C
4

Splitter (iC
4
H
10



nC
4
H
10
)


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 c
ondensed 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 reboil
er system
. The s
plitter 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 n
a
phtha
from the column bottoms is fe
d to the Naphtha Hydrotreater section and subsequently
the Catalytic Reformer feedstock.




The t
hermodynamics

suitable for simulating these hydrocarbon mixtures are the equation of states
Soave


Redlich
-

Kwong (SRK) fo
r pressures >1 bar and Peng Robinson for pressures >10 bar.



17


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 predictio
n, 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

Componen
t

Overhead

Bottoms

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

H
2
S

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.




18



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
t
ray specification
as follows:


Key Tray Specifications

Column Height

m

51.8

Downcomer Area (cen
tre)


m
2

0.86

Column Diameter

m

2.9

Tray Spacing


m

0.6

Number of Trays

no

74

Hole Diameter


mm

39

Weir Length (side)

m

1.859

Total Hole Area


m
2

0.922

Weir Length (centre)

m

2.885

Outlet Weir Height


mm

51

Liquid Flowpath Length

m/pass

0.967

Tray Thickness


mm

2

Active Area

m
2

4.9

Number of Valves


no/tray

772

Downcomer Area (side)

m
2

0.86

Free Fractional Hole Area

%

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.












The total hole area is shown as 0.922m
2

which
is in ratio to the active area of 4.9m
2

giving the
free fractional hole area of 18.82%.


The simulation is now shows the following
results:





1.859

0.335

Area 0.43

2 Passes


19


Section 6

Hydrotreater


Hydrotreater Simulation

Case/File Name

Description

R6.01

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, 1
st

Edition, March 2000.

3.

W.L.Nelson, “Petroleum Refinery Engineering”, 4
th

Edition, McGraw Hill, 1958.


Process
(1
, 2
)


H
ydrotreater
s are used

to selectively remove u
ndesirable elements and

hydrogen
saturate unsaturated
co
mponents.


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.


H
2

reacts with mercaptans (1H
2
), thiophenes (3H
2
) and benzothiphenes (5H
2
) produce H
2
S

H
2

reacts with pyridine (5H
2
) produces NH
3

H
2

reacts with phenols (1H
2
) produces H
2
O


Hydrotreater Hydrogen Usage and Losses

(2)

Units

scf/barrel fresh feed

sm
3
/m
3

fresh

Remarks

Reactions(basis fresh feed

100

=
㔰M
=
ㄸN
J

=
啮r慴ar慴敤⁣潭灯湥n瑳‾e
2

Solubility(basis fresh feed)

10
-

20

1.8

=
㌮P
=
=
m畲来
=
㐰4
J
=
㄰M
=
㜮㈠
J
=

=
䑥灥湤⁈
2

in makeup gas

Recycle

500
-

1500

89
-

267

Maintain H
2
/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
an
d 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.







20



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.















21



Section 7

Catalytic
Reformer


Reformer Simulation

Case/File Name

Description

R7.01

Catalytic Reformer


References

1.

G.L.Kaes, “Refinery Process Modelling”, Athens Printing Company, 1
st

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
use
s

three
fixed bed reactors in series.


E
ndothermic dehydr
ogenation
reactions take place in
the first

two reactors requiring fired inter
-
heaters to raise the

temper
ature for the
following reactor with h
ydrocra
cking reactions being significant in the final reactor.



Reforming catalysts are subje
ct to poisoning by hydrogen sulph
ide and other sulfur compounds,
nitrog
en, and oxygen which are removed from the naph
tha by mild hydrotreating.


The primary reformer
feed stock is virgin
(uncracked) naphtha f
rom the crude distillation process and

other naphtha stocks of
suitable

boiling point range are accept
able after hydrotreating.


T
he reactions do not occur evenly th
rough 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 ou
tlet temperature.
U
sing 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.


P
re
-
treated naphtha is
combined wi
th 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

Reactor 1

Reactor 2

Reactor 3

Inlet temperature

°F

937

937

937

Inlet pressure psia

413

394

394

Measures ∆T (Typical) °F
=
㘰
㤰
J
N㌰F
=
㌵
㐰
J
S〩
=


(1)

(10
-
20)

Recycle MMSCFD

0.5
-
1% Naphtha Feed



Catalyst Volume ft
3

274

411

910


Note 1 Simulation temperature drop is
much larger due to the example reactions considered

T
he temper
ature 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 pressu
re 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

Separator

Temperature °F

100

Pressure psia

290

H2 Purity

0.79



22



Stabilizer parameters


T
he
liquid is feed to the stabilizer to remove the light ends.


Reformer stabilizers generally have 30 to 36
actual trays

with

o
verall tray
efficienc
ies 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 recov
ery plant and t
he column
bottom
s product is
stabilized reformate.




Operating Data

Stabilizer

Number of
stages

36

Feed tray

19

Feed temperature °F

297

Tray 1 temperature °F

257

Bottom temperature °F

446 (Simulation 488)

Partial condenser pressure psia

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 produ
cts 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
-
Hinsh
elwood,
Arrhenius and Smith. The Smith model considers the following groupings:


N
aphthene
+ H
2




P
araffin


Naphthene
s





A
romatics
+
3H2

Hydrocracking of paraffin

Hydrocracking of naphthenes


Pseudocomponent groupings will include specific boiling rang
es such as for C6 to C11 paraffins, C6 to
C11 naphthenes, benzene, toluene and C8 to C11 aromatics.











23



Case 7.01 Catalytic Reformer


Flowsheet is shown:




Thermodynamics selected:

K
-
Peng Robinson

H
-
Peng Robinson


The configuration for catalytic Reactor 3 is shown.







24



Section
8


Amine Treatment


Vacuum Still Simulations

Case/File Name

Description

R8.01

Sour Gas Treatment


References

1.

A.Kohl and R.Nielsen, “Gas Purification”, Gulf Publishing , 5
th

Edition, 1997


Proces
s



Chemical absorption of CO
2

and H
2
S with amines provides the most cost effective mean
s of obtaining
high purity vapo
r from sour gases in a single step.

The process is well established for refinery gas
sweetening which are carried o
ut 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 v
ary 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 ten
ds to
degrade as the temperature rises, increasing the replacement expense and involving the removal of
degradation products.


MEA has
a substantially higher vapo
r pressure than other amines and a water wash at the top of the
absorber can be used to minim
ize 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
stripp
ed 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 condensa
tion of an immiscible
hydrocarbon liquid which will promote foaming in the absorber.




25


Section 9

Miscellaneous

Application


Miscellaneous Application Simulations

Case/File Name

Description

R9.01

Biodiesel Blending


Biodiesel
Blending


Diesel is blended
with Methyl Ester

(ME) during
ship offload
ing. 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
disch
arge flow can vary from

800
m
3
/h
to
1000 m
3
/h
at a maximum pres
sure of 10 barg.


The mass balance on the streams give:







Where:


D
E


Diesel in Methyl Ester Flow

(
m
3
/h
)

E

Methyl Ester Flow


(
m
3
/h
)

V
E

Methyl Ester Volume Fraction

D
S

Diesel Flow from Ship


(
m
3
/h
)

V
P

Bio
-
diesel Product Volume Fraction

We have:


D
E
E
V
E
E



and

D
D
E
E
V
S
E
P




Rearranging gives:




V
V
1
E
D
E
E
E




and substituting for
D
E

leads to the following:



V
V
1
D
V
E
E
P
S
P



and



1
V
V
1
D
E
D
P
E
S
E





Ester Blend to Ship Flow Ratios

Methyl Ester

Product Blend

V
E

/ V
P


Ship to Ester
Blend

Flow Ratio

%

%

100

15

6.667

0.176


10

10.0

0.111


5

20.0

0.053

80

15

5.33

0.231


10

8.0

0.143


5

16.0

0.067


A process control system would set t
he

blender flow ratios by entering the ME blend
(V
E
) and Final
Product blend
(V
P
). The ME blend flow (E+D
E
) required for a “wild” Ship Discharge flow (D
S
) is
calculate
d. The flow ratio controller would

mani
p
ulate 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.






D
E

E


D
S

V
E

V
P

D
E

+ E


D
S

+ D
E

+ E



26



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 Physical Property Data at 15ºC

Fluid

Density

Viscosity

kg/m
3

cps

Diesel

807.15

2.78

Methyl Ester !00%

876.43

7.48

Methyl Ester 80%

864.74

5.72




The controllers are configured as shown:




The simulation
results are in accordance with the theory developed above.



The maximum and minimum es
ter blend flowrates obtained are
230.7 and 42.1 m
3
/h.








27


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 ft
3

atm/ R lb mol


API gravity formula
e


API gravity

(API ρ)

of petroleum liquids is determined from specific gravity (SG)

at 60°F:




The specific gravity of petroleum liquids can be derived from the API gravity
:





Heavy

oil with a specific gravity of 1.
0 (density
water at 60°F) has

an API gravity of:




Crude oil is
often measured in metri
c tons (1000kg). The
number of barrels per metric ton

for a given
cru
de oil based on its API gravity is calculated from:




Where 1 bbl = 0.159 m
3

A

metric ton of West Texas Intermediate

39.6° API would contain about 7.6 barrels.

28



Reid Vapor Pressure (RVP)

A measure
of gasoline

volatility being

defined as the absolute vapor pres
sure 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 p
etroleum 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


D
efined 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 c
ompounds
.


VABP and MeABP

Petroleum fractions are cuts with specific boiling point ranges, API gravity and viscosity. Each cut can
be divided int
o 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 po
int (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
,

7
0 and 90% distilled:




MeABP is calculated
from
:


Where ∆ is given by
:






MeABP (°R) is used in the definition of the Watson K which is given by
:




Factors
Note 1

Prefix

Symbol

10
-
12

E
-
12

pico

p

10
-
9

E
-
09

nano

n

10
-
6

E
-
06

micro

μ
=

J
3

E
-
03

milli

m

10
-
2

E
-
02

centi

c

10
-
1

E
-
01

deci

d

10
1

E01

deca

da

10
2

E02

hecto

h

10
3

E03

kilo

k

10
6

E06

mega


M
Note 2

10
9

E09

giga

G

10
12

E12

tera

T


Note 1

Tip for setting power, make equal to number 0’s so 0.00001 = 10
-
5

and 100000 = 10
5

Note 2 Refinery industry practice sometimes uses MM
to signify 10
6


29




30



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 (Refer
ence Crane 410M)

Service Conditions

Fluid

Reasonable Velocities

Pressure Drop

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


General Service

Liquids pumped

Non viscous

1.0 to 3.0

3.2 to 10

0.05

Heating Short Lines

Saturated Steam

0 to 1.7 bar

20 to 30

65 to 100


Process piping

Saturated Steam

>1.7 bar

30 to 60

100 to 200


Boiler and turbine leads

Superheated Steam

14 and up

30 to 100

100 to 325


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

d in (d/6 + 1.3) ft/s

d mm (d/500 + 0.4) m/s

Pump discharge line for

d in (d/3 + 5) ft/s

d mm (d/250 + 1.5) m/s

Steam or gas

d

in 20d ft/s

d mm 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


(1.3 + d/6) ft/s


0.4 psi /100 ft

Liquid Pump discha
rge


(5.0 + d/3) ft/s


2.0 psi /
100 ft

Steam or gas



(20d) ft/s


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.


31



Process Simulation

Procedures and Convergence


Steady state simulation proves

the capability to achieve stable and reproduci
ble operating conditions
with acceptable product purity, yield and

cycle times to satisfy the commercial requirements and the
safety and environmental issues for the regulatory au
thorities.


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.







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 p
oints
.

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.

F
or refinery operations the Tower UnitOp is more suitable as it includes pump around and stripping
facilities.

The s
tage 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 reboil
er 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 sta
ge 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
sp
ecification,
mass fraction of a key componen
t 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’ th
e
specifications toward the
target specification
.
Use the following procedure:


Feeds

Recycles

Products


32



1.

Set up the column
: number of stages, condenser, reboiler, operating pressure.

2.

Generate TPxy and RCM plots to v
erify that the target is thermo
dynamically feasible with the
selected VLE K model.

3.

O
n the SPECIFICATIONS page, set ‘loose’ specificatio
ns such as ‘Reflux Ratio’ and ‘

Bottoms F
lowrate
’ or Reboiler Heat Input.

4.

Run the column and converge. Change the specifications if necessary.

5.

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.

6.

On the SPECIFICATIONS page change to more tight specifications. Run the c
olumn.


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.


Of
ten
,

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.

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.

2.

Remove non

key components from the feed(s) to obtain the first convergence.
Now s
et the
in
i
tial flag to
0 Reload Column Profile
,

return the other components, and run the unit again.

3.

Specify a larger number of iterations on the convergence page of the column dial
og
. The
default is 50, but possibly

52 iterations will find
the answer.

4.

Try an alternate column model. If you are currently using the SCDS try the same separation
with a TOW
E
R or vice vers
a. 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.

5.

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 withou
t
cryogenic
temperatures. Changing condenser mode to partial allows the light ends gases to

slip
past the condenser.