Acetic Acid Process Plant Design

1
Acetic acid process plant design
By
Hisham Albaroudi
Karen Atayi
Cristian Baleca
Enoch Osae
Sinthujan Pushpakaran
Alexander Taylor
School of Chemical Engineering
Faculty of Science and Engineering
University of Hull
June 2016

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Table of Contents
1.) Executive Summary 1
2.) Process Selection 2
2.1) Process Technology Selection 2
2.2) Process Flowsheet Development 5
2.3) Process Description 6
3.) Piping and Instrumentation Diagram (P&ID) 7
4.) Mechanical Design of Unit Operations 8
4.1) Reactor 8
4.1.1) Introduction 8
4.1.1.1) Propionic acid 9
4.1.1.2) Water – Gas shift 9
4.1.1.3) Methyl acetate 9
4.1.1.4) Methyl iodide 9
4.1.2) Reactor P&ID 10
4.1.3) Design Method 11
4.1.3.1) Reactants and product 11
4.1.3.2) Variations 12
4.1.3.3) Energy Balance and Heat of reaction 12
4.1.3.4) Choice of reactor 13
4.1.4) Reactor Specification 13
4.1.4.1) Choice of material 14
4.1.4.2) Vessel support 14
4.1.4.3) Piping sizing 14
4.1.4.4) Nozzles 15
4.1.4.5) Heat dissipation and vessel insulation 15
4.1.4.6) Shut down 16

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4.1.4.7) Process safety 16
4.1.4.8) Process control 16
4.1.5) Agitator Specification 17
4.1.6) Conclusion 19
4.1.7) Engineering Drawing of reactor 21
4.2) Flash Tank 22
4.4.2) Flash Tank P&ID 23
4.2.4) Flash Tank Specification 25
4.2.4.1) Process controls and safety 27
4.1.6) Engineering Drawing of Flash Tank 29
4.3) Drying Distillation Column 30
4.3.2) Drying Distillation Colum P&ID 31
4.3.4) Drying Distillation Column Specification 32
4.1.7) Engineering Drawing of Drying Distillation Column 34
4.4) Heavy Ends Distillation column 35
4.4.2) Heavy Ends Distillation Colum P&ID 37
4.4.3) Acetic Acid Properties within the Column 38
4.4.4) Propionic Acid Properties within the Column 38
4.4.5) Relative Volatility 39
4.4.6) Heavy – Ends Distillation Column Specification 39
4.4.7) Summary of Design Data 41

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4.1.9) Engineering drawing of Heavy Ends Distillation Column 42
4.5) Absorption Column 43
4.5.2) Absorption Column P&ID 44
4.5.4) Absorption Column Specification 45
4.5.4.1) Design Considerations to Account for Drawback of Unit 45
4.5.4.2) Choice of Packing 46
4.5.4.3) Choice of Absorption Tower Equipment 46
4.5.4.4) Materials of Construction 47
4.5.4.5) External Equipment 47
4.5.4.6) Safety Control 48
4.5.4.7) Safety Considerations 48
4.5.6) Engineering Drawing of Absorption Column 50
4.6) Storage tank – Acetic acid 51
4.6.2) Storage Tank – Acetic Acid P&ID 52
4.5.4) Storage Tank Specification 53
4.5.6) Engineering drawing of Storage Tank – Acetic Acid 56
5.) Process Control and Instrumentation 57
5.1) Introduction to Process Control and Instrumentation 57
5.2) Objectives of Process Control 57
5.3) Implementation of Control Systems in our Design 58

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5.3.1) Distillation Column Control 59
5.3.2) Reactor and Flash Tank Control 61
6.) Process Economics 62
6.1) Market Analysis 62
6.2) Costing 63
6.2.1) Feedstock price estimation 64
6.2.2) Capital Cost Estimation 65
6.2.2.1) ISBL 65
6.2.2.2) Installation Factor 66
6.2.2.3) OSBL 66
6.2.2.4) Engineering costs 66
6.2.2.5) Contingency costs 66
6.2.2.6) Fixed Capital Investment 67
6.2.3) Working Capital 67
6.2.4) Total Investment 67
6.2.5) Operating Expenditure 67
6.2.6) Revenue 67
6.2.7) Gross Profit 67
6.3) Project Financing 68
6.3.1) Financing Bank Loan 68
6.3.2) Financing Investment 70
6.3.3) Net Profit 71
6.3.4) Cumulative Cash Position 72
6.3.5) Return of Investment 73
7.) Process Safety 74
7.1) Safety legislations 75
7.2) Hazard Identification 75

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7.2.1) Material Hazard 75
7.2.2) Material Toxicity 78
7.2.3) Flammability 79
7.3) Operating conditions hazard 80
7.3.1) Pressure relief strategy 80
7.3.2) High pressure response measures 81
7.3.3) Fire prevention strategy 81
7.3.4) Fire and gas detection 82
7.3.5) Noise 83
7.3.6) Loss of containment 83
7.4) Emergency Response plans 83
7.4.1) Fire 83
7.4.2) Explosion 85
7.4.3) Overpressure 85
7.4.4) Toxic release 85
7.4.5) Flooding 85
7.4.6) Earthquakes 86
7.4.7) Human error 86
7.4.8) Personal Protection Equipment 86
7.5) HAZOP 86
7.5.1) Scope of work 87
7.5.2) Term of Reference 87
7.5.3) Team Membership 88
7.5.4) Safety conclusions 89
7.5.5) Marked up P&ID 90
7.5.5) HAZOP findings 91
8.) Environmental Protection 95

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8.1) Process Selection 95
8.2) Plant Location – Environmental Considerations 95
8.3) Noise 96
8.4) Odour 96
8.5) Traffic 96
8.6) Catalyst and Water Requirement 97
8.4) Methanol Feed 97
8.5) Energy Recovery 98
8.6) Storage & Handling of Raw Materials and Product 98
8.6.1) Carbon Monoxide 98
8.6.2) Methanol 98
8.6.3) Acetic Acid 99
8.7) Undesired products: By- and Co- products 99
8.7.1) Propionic Acid 100
8.7.2) Carbon Dioxide and Hydrogen 100
8.7.3) Methyl Iodide 101
8.7.4) Aqueous and Organic Discharges 101
9.) Plant Layout and Location 102
9.1) Plant location 102
9.2) Plant layout 104
9.2.1) Site Flow Plan 107
10.) Appendices
APPENDIX [A] – Minutes 108
Meeting Week 13 – 9th
May 2016 109
Meeting Week 12 – 3rd
May 2016 110
Meeting Week 11 – 22nd
April 2016 111
April 2016 112

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April 2016 113
April 2016 114
March 2016 115
Meeting Week 6 – 29h
February 2016 116
Meeting Week 5 – 22nd
February 2016 117
February 2016 118
February 2016 119
Meeting Week 2– 5th
December 2015 120
November 2015 121
APPENDIX [B] – Reactor Calculations 121
APPENDIX [C] – Flash Drum Calculations 140
APPENDIX [D] – Calculations for Drying Distillation Column 151
APPENDIX [E] – Calculations for Heavy – Ends Distillation Column 160
APPENDIX [F] – Calculations for Absorption Column 177
APPENDIX [G] – Calculations for Acetic Acid Storage Tank 189

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1. Executive Summary
The purpose of this document is to present a potential design to the client to build an acetic acid
(CH3COOH) plant in the United Kingdom. The plant will have the capacity to produce 400,000 tonnes per
annum of acetic acid base product from a feedstock of methanol and carbon monoxide. As an overview,
the methanol carbonylation process is highly efficient in that it produces acetic acid with more sought after
selectivity and purity. Although, the oxidation of ethylene is more environmentally friendly, it can only be
operated for a capacity of up to 200,000 tonnes per year, while the oxidation of hydrocarbons route for
acetic acid production is cheaper to run but it does not produce pure acetic acid and greatly affects the
environment as a result of its CO2 emissions. Even though the oxidation of ethylene and methanol
carbonylation processes do not pose much threat to the environment, the latter is still more environmentally
friendly as it produces less waste and recycles most of its reactants.
Environmental Impact Assessment has been proven successful in outlining the main environmental issues
in relation to this project. The general location considerations linked to the potential pollution produced
(odours, noise, traffic) has been analysed, justifying the measures that will be put in place to minimize
them. The handling of raw materials and the final product both on and off site has been studied in depth in
order to outline the features and add-ups that can be applied to reduce the impact on the environment:
such measures are mainly related to the close monitoring and the implementation of safety measures to be
applied whenever a containment vessel were to mechanically fail for any reason: it has been concluded
that appropriate containment chambers and process control and instrumentation are the significant routes
to apply. An eco-friendly engineering strategy has been applied to this project when dealing with significant
by-products. Although the generation of highly corrosive chemicals (such as methyl-iodide) has not been
possible to be prevent, other “waste” compounds produced in the system have been proved to be
commercially useful (propionic acid). Certainly, a comprehensive recycling system within the plant (and
especially to the reactor) is a successful strategy to ensure that non desired products are dealt with the
purpose of minimizing waste. Furthermore, even when waste streams are unable to be recycled and re-
used in the system, appropriate techniques to dispose of them have been developed in compliance with
environmental regulations and ethical considerations i.e. flares, aqueous discharge basin and pipeline.
Careful consideration of all the hazards present on the plant are outlined in the following report which
highlights efficient ways of maintaining a safe environment for the production of acetic acid.
In addition to environmental methodologies, principles of process control and instrumentation have been
applied throughout the design stage of this project with the aim of creating a process that is ultimately safe,
that complies with all the necessary safety regulations, efficient, that will not suffer unnecessary downtime to
avoidable failures and maintenance being carried out on key piece of process equipment and not suffer
performance impairments due to poor design, as well as being economically stable, linked to the plants
efficiency, an efficient plant will bring a certain amount of economic stability in addition to ensuring
unnecessary equipment or instrumentation is not put in place.

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Based on market research it is possible to conclude that the acetic acid market is projected to rise the
upcoming years on a global basis. Economic evaluation of this project indicates viability, the return of
investment is 53% and the net profit of £1,378,000,000 is very lucrative figure for a 20-year investment.
Both a bank loan and private equity investments would generate a greatly positive profit, although a bank
loan would represent a significantly more profitable route. The project payback time of 2 years
demonstrates that this project is highly feasible and has the potential to attract numerous investors.
2. Process Selection
2.1 Process Technology Selection
The methanol carbonylation, direct oxidation of n-butane and direct oxidation of ethylene are the three most
widely implemented methods to manufacture acetic acid.
Methanol carbonylation also known as the “Monsanto process” was initially developed by BASF in 1960.
The process operates at 180 – 220 o
C and 30 – 40 atm via the use of a rhodium catalyst, leading to energy
costs set to a bare minimum. Although it operates at such low operating conditions, the process provides
with high selectivity of acetic acid (Yoneda et al., 2001). The final product holds great purity due to the
selectivities of methanol and carbon monoxide, which are 99% and 90% respectively (Yoneda et al., 2001).
The process outlines continuous supply of methanol and carbon monoxide into the reactor. The
combination of exhaust gas produced from the reactor and purification section are recovered as light-ends
and recycled back into the reactor. Consequently, the acetic acid produced from the reactor is separated as
a side-cut and delivered to the dehydration column (Sano et al., 1999). Acetic acid and water mixture are
then released at the top of the column and back to the reactor while propionic acid is taken to the
subsequent column. Further purification takes place and acetic acid is generated as a side-cut. Continuous
recycling of overhead and bottoms found in fractional column into reactor take place (Sano et al., 1999).
The main raw materials for this process are methanol and carbon monoxide. In the reaction process,
methyl iodide is added to the rhodium complex, which consecutively migrates to a carbonyl group and
reacts with CO to form the rhodium-acetyl complex (Kinnunen and Laasonen, 2001). The excess water
readily hydrolyses the acetyl iodide (CH3COI) to produce acetic acid and hydrogen iodide in order to
complete the catalytic cycle (Yoneda et al., 2001). However, a quantity of water (14 – 15 wt.%) is required
in order to maintain stability and activity of the catalyst, thus separation of water from acetic acid requires
excessive amount of energy, further limiting storage capacity (Wittcoff et al., 2013). Methanol carbonylation
produces propionic acid as the major by-product of this process (Sunley and Watson, 2000), present as an
impurity in methanol feed (Yoneda et al., 2001). In order to lower the yield of propionic acid produced, it is
suggested to decrease the amount of acetaldehyde produced by the rhodium catalyst (Yoneda et al.,
2001).
The direct oxidation of hydrocarbons route occurs through pumping of an ethane and oxygen mixture at
515 K and 16 bar in a multi-tubular reactor (Smejkal et al., 2005). The product formed is cooled to 303 K via

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two steps, initially through formation of high-density steam and subsequent separation of formed gas and
liquid mixture in a flash (Smejkal et al., 2005). The acetic acid-water mixture produced is then separated in
a rectification column and pure acetic acid is generated as the bottom product. The resulting gaseous
stream, made up of ethane, ethylene and CO2, is recycled back into the system. CO2 is separated into an
absorber, while ethylene and ethane are put back into the feedgas (Soliman et al., 2012). Nonetheless, the
oxidation of n-butane requires large amounts of water and generates a dilute acetic acid solution of which
concentration is highly energy intensive, as a result the yield of acetic acid produced is lower than the one
obtained in other processes (Sano et al., 1999). A vast amount of by-products are formed, some of which
are propionic and formic acids (Riegel, 2007). Furthermore, this particular process requires large quantities
of water, hence water gas shift reaction is a major drawback as extensive CO2 is produced as a result
(Wittcoff et al., 2013). In essence, the oxidation of hydrocarbons process is cheaper to run as a result of its
feedstock, but at the cost of being less efficient as it produces more waste and a lower grade chemical.
Flexibility in the process allows it to produce a purer acetic acid with high selectivity, however extensive
operation expenses are necessary.
The production of acetic acid through direct oxidation of ethylene was first proposed by Showa Denko K.K..
The process occurs through the mixture of ethylene and oxygen in their vapour phases at 160 – 210 o
C
over a solid catalyst (Xu et al., 2010), the acetic acid generated is of high selectivity. The reaction is
initiated from the cooling of gas produced in the reactor to ambient temperature, where the products of
acetic acid, water and other organic compounds are condensed and separated. The condensate transfers
to the crude acetic acid tank, while the compressor pressurizes the un-condensed gas back to the reactor.
Light-end products such as acetaldehyde, ethyl acetate and ethanol are removed through distillation,
allowing acetic acid and light-ends compounds to migrate to the purification section where pure acetic acid
is produced (Sano et al., 1999). This process produces large amounts of heat which is recovered as steam
and used in the purification section as a source of heat (Sano et al., 1999).The process meets the
requirements of being both competitive and environmentally friendly. Although it rivals methanol
carbonylation, the process is only efficient with smaller plants of about (100-250 kt/a) (Sano et al., 1999)
and considering the fact that the feedstock price of ethylene is more expensive than raw materials used in
the other processes mentioned, thus economically it will not be as profitable as the methanol carbonylation
process given that the selectivity of both is of 90%.
Overall, the methanol carbonylation process is highly efficient in that it produces acetic acid with more
sought after selectivity and purity. The oxidation of ethylene is more environmentally friendly, however it
can only be operated for a capacity of up to 200,000 tonnes per year and the oxidation of hydrocarbons is
cheaper to run but does not produce a pure acetic acid product. The oxidation of hydrocarbons highly
affects the environment as a result of its emissions of CO2, whilst both the ethylene oxidation and methanol
carbonylation don’t pose much threat to the environment, the latter is still more environmentally friendly as
it produces less waste and recycles most of its reactants. This enables the process to be continuous and
thus economically beneficial.

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2.2 Process Flowsheet Development
Reactor
1
2
3
Flash tank
Light ends
distillation column
Drying
distillation column
Heavy ends
distillation column
8 12
Scrubber
9
113
14
1074
8
1
2
3
6
9
17
16
11
15
Pressure reduction
valve
5
Figure [1]: Process Flow Diagram and Material balance of process.
Steam 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17
Molar Flowrate
[kg/hr]
Methanol 0.00 27044.00 0.00 56.15 56.15 1.15 56.07 0.00 56.07 Trace 0.00 0.00 0.00 0.00 0.00 Trace 57.29
Carbon monoxide 25209.36 0.00 0.00 207.44 207.44 200.88 207.55 0.00 207.55 0.00 0.00 0.00 0.00 0.00 0.00 3.02E+02 105.97
Acetic acid 0.00 0.00 0.00 98543.49 98543.49 611.25 88689.05 9854.21 5918.53 8.28E+04 29024.25 53676.293 53491.87 184.35 2674.63 1.60E+02 9045.00
Water 0.00 0.00 720.40 234.38 234.38 3.19 210.90 23.41 210.90 Trace 0.00 0.00 0.00 0.00 0.00 Trace 214.13
Ethanol 0.00 273.14 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Trace 0.00 0.00 0.00 0.00 0.00 Trace 0.00
Propionic acid 0.00 0.00 0.00 487.99 487.99 2.02 439.24 48.89 0.00 4.39E+02 0.01 439.24 0.01 439.24 0.00 Trace 2.02
Carbon dioxide 0.00 0.00 0.00 6256.54 6256.54 947.56 6256.36 0.00 6256.36 0.00 0.00 0.00 0.00 0.00 0.00 2.43E+03 4773.63
Hydrogen 0.00 0.00 0.00 8.69 8.69 72.79 8.69 0.00 8.69 0.00 0.00 0.00 0.00 0.00 0.00 7.99E+01 1.60
Methyl acetate 0.00 0.00 0.00 1243.25 1243.25 24.77 1242.89 0.00 1242.89 Trace 0.00 0.00 0.00 0.00 0.00 Trace 1268.02
Hydrogen iodide 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 Trace 0.00
Methyl iodide 0.00 0.00 0.00 2298.05 2298.05 60.51 1839.02 459.76 1839.02 0.00 0.00 0.00 0.00 0.00 0.00 Trace 1898.95
Total 25209.36 27044.00 720.40 109343.00 109343.00 1924.24 98955.71 10386.75 15742.53 83213.17 29024.26 54117.67 53494.16 623.51 2674.71 2972.623 17368.82
Temperature
[o
C]
25 25.53 26.87 160 104.45 160 104.45 104.45 79.88 117.65 141.66 117.69 117.58 131.41 117.58 18.31 69.71
Pressure
[bar]
31 31 31 30.40 1 30.40 1 1 1 1 2 1 1 1 1 1 59

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2.3 Process Description
Aspiring Consulting have decided to manufacture acetic acid using the Monsanto process. The first stage
of the process involves the injection of the raw materials methanol, carbon monoxide and water into the
reactor to initiate the methanol carbonylation process. The carbonylation reaction is carried out in a stirred
tank reactor on a continuous basis. Liquid is then removed from the reactor through a pressure reduction
valve. This then enters the flash tank, where light components of methyl acetate, methyl iodide, some water
and product acid are removed as vapour through the top of the vessel. The gases are fed forward to the
distillation train for further purification whilst the remaining liquid in the flash tank, containing the dissolved
catalyst, is recycled to the reactor.
Liquid from the reactor enters the lower half of a multiple-tray distillation column operating at conditions
above atmospheric conditions. Hydrogen iodide present in the feed stream is concentrated in the acetic
acid solution in the bottom of the column. This stream is recycled back to the reactor. Carbon monoxide,
water, methyl iodide and some entrained hydrogen iodide comprise the overhead stream from the column,
which passes through a condenser and phase separator where uncondensed gas is directed towards the
scrubber.
The condensate separates into two phases: a water phase consisting of some organic compounds and an
organic phase (methyl iodide) containing some water. The organic phase is recycled to the reactor whilst
part of the water phase is used as reflux in the distillation column and excess is recycled to the reactor.
Solution of acetic acid in water containing some iodide, catalyst and by-products is withdrawn from the
bottom of the column and introduced into a second multiple-tray distillation column operating at conditions
above atmospheric conditions. In this column, water and remaining inerts are withdrawn overhead and
directed towards to the scrubber. A portion of the condensate is returned as reflux to the column and
excess is recycled to the reactor. To avoid accumulation of water in the system, a portion of the water
separated in the column is discarded.
Residual hydrogen iodide in the feed stream to the column concentrates at a location near the middle of the
distillation column. By continually withdrawing the solution containing hydrogen iodide from the middle of
the distillation column, virtually all of the hydrogen iodide is removed from the column. This solution can be
recycled directly to the reactor or alternatively to the lower half of the previous distillation column, where it is
concentrated and removed with the bottoms stream of that column.
Acetic acid product is withdrawn from the drying column without further processing; acetic acid vapour is
withdrawn from the top of the column and passed through a condenser from which it is pumped to storage.
Liquid acetic acid containing residual catalyst is periodically withdrawn from the top of the column and
recycled to the reactor.

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3. Piping and Instrumentation Diagram (P&ID)
DC-501
P-501
P-502
V-502
V-503
V-504
V-505
FI
FIFI C
FI C
V-501
LI C
LI
2" 304SS
V506
TI
FI C
36" 304SS
V-507
TI FI C
DC-401
P-401
P-402
8" 304SS
V-402
V-403
V-404
V-405
FI
FIFI C
FI C
V-401
LI C
LI
4" 304SS
V-406
TI
FI C
V-407
TI FI C
6" 304SS
DC-301
P-301
P-302
V-302
V-303
V-304
V-305
FI
FIFI C
FI C
V-301
LI C
LI 10" 304SS
V-306
TI
FI C
TI FI C
V-307
Acetic acid
Propionic ac id
R-201
V-210
V-209
V-208
V-207
FI
FI FI C
FI C
S-601
Fl are
6" HC
V-201
LI
FI C
V-202
V-203
V-205
V-204
PIFI C
F-201
V-206
FI C
LI
V-211
FI FI C
6" HC
80" CS
V-308
FIFI C
V-212
FIFI C
CH3OHfeed
H2Ofeed
COfeed
P-101 P-102
V-102V-101
V-104V-103
FI
FIFIC
FIC
P-103 P-104
V-106V-105
V-108V-107
FI
FIFIC
FIC
1" CS
V-110V-109
V-112V-111
FI
FIFIC
FIC
3" CS
P-602
P-601
V-617
V-619
V-620
FI
FIFI C
FI C
12" 304SS
V-602 V-604
V-411
V-410
V-409
V-408
FI
FI FI C
FI C
12" HC
V-601 V-603
V-609
V-610
V-612
V-611
V-511
V-510
V-509
V-508
FI
FI FI C
FI C
V-618
V-613
V-614
V-616
V-615
LIFI C
V-605 V-607
V-606 V-608
P-503
P-504
LAH
LAL
LAH
LAL
LAH
LAL
LAH
LAL
Process
Steam
Process
Steam
Process
Steam
Process
Water
Process
Water
Process
Water
V-508
FI C FI
V-408
FI C FI
V-308
FI C FI
Figure [2]: P&ID showing all unit operation.
Tag Number R-201 Tag Number F-201 Tag Number DC-301 Tag Number DC-401 Tag Number DC-501 Tag Number S-601 Aspiring Consulting LTD. KEY
Service Reactor Service Flask Tank Service
Light Ends Distillation
Column
Service
Drying Distillation
Column
Service
Heavy Ends Distillation
Column
Service Scrubber
Process
P&ID
Client
UoH
plc
CS Carbon Steel
Design Press 33 bara Design Press 1.1 bara Design Press N/A Design Press 1.1 bara
Design
Press
1.52 bara Design Press 1.5 bara HC Hastelloy C
Design Temp 176 o
C Design Temp 115 o
C Design Temp N/A Design Temp 125 o
C
Design
Temp
150 o
C Design Temp 50 o
C Location Immingham 304SS
304 Stainless
Steel
Height 8.95 m Height 10 m Height N/A Height 23.85 m Height 21.6 m Height 10 m

15
4. Mechanical Design of Unit Operations
4.1 Reactor
4.1.1 Introduction
The overall reaction for the production of acetic acid is given by the kinetic equation (Cheng and Kung,
1994):
𝐶𝐻3 𝑂𝐻 + 𝐶𝑂 → 𝐶𝐻3 𝐶𝑂𝑂𝐻
ΔG = -72.79 kJ/mol ΔH = -133.82 kJ/mol
The enthalpy values collected in the literature clearly indicate the exothermic nature of the process; energy
is being produced within the reactor. Typical operating conditions emphasized in literature for the Monsanto
process lay in the following ranges (Cheng and Kung, 1994):
Pressure Temperature
30-60 bar 150-200 oC
Table [4.1] – Ranges for operating conditions.
In order to model the simulation, different values for pressure and temperature have been selected for the
reactor; the values that appeared to give the highest selectivity yet maintaining a relatively lower generation
of by-products are 160 o
C and 30 bar respectively.
The process uses a Rhodium and Iodine complex to catalyse and promote the reaction, and this catalyst is
highly selective especially in relation to methanol conversion. Due to the high activation energy, the
reaction would not occur without the aid of a catalyst.
The catalytic reactions are listed below, and are all equilibrium limited for the purpose of the reaction
simulation. This is relatable to the fact that catalysts and promoters do not affect the stoichiometry and heat
of reaction:
CH3OH + HI ↔ CH3I + H2O
CH3I + CO ↔ CH3COI
𝐶𝐻3 𝐶𝑂𝐼 + 𝐻2 𝑂 ↔ 𝐶𝐻3 𝐶𝑂𝑂𝐻 + 𝐻𝐼
As far as the process simulation is concerned (Aspen PLUS), the main methanol carbonylation reaction is
kinetic, and some values in the literature have been investigated in order to model the simulation.
The following information has been used to model Aspen Plus simulation:
𝑇𝑎𝑟𝑔𝑒𝑡 𝑝𝑟𝑜𝑑𝑢𝑐𝑡𝑖𝑜𝑛 𝑜𝑓 𝑎𝑐𝑒𝑡𝑖𝑐 𝑎𝑐𝑖𝑑 = 400,000 𝑡𝑜𝑛𝑠 𝑓𝑜𝑟 8000 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑛𝑔 ℎ𝑜𝑢𝑟𝑠 (𝑎𝑠 𝑝𝑒𝑟 𝑑𝑒𝑠𝑖𝑔𝑛 𝑠𝑝𝑒𝑐𝑖𝑓𝑖𝑐𝑎𝑡𝑖𝑜𝑛)
= 50,000 𝑘𝑔/ℎ𝑟

16
In accordance to the stoichiometry and molecular weight of the reactants, the required input methanol and
CO have been calculated for this purpose:
𝐶𝑂 →
186,666 𝑡𝑜𝑛𝑠
0.9
= 207,407
𝑡𝑜𝑛𝑠
𝑦𝑒𝑎𝑟
= 25,925 𝑘𝑔/ℎ𝑟
𝑀𝑒𝑡ℎ𝑎𝑛𝑜𝑙 →
213,333 𝑡𝑜𝑛𝑠
0.99
= 215,488
𝑡𝑜𝑛𝑠
𝑦𝑒𝑎𝑟
= 26,936 𝑘𝑔/ℎ𝑟
Due to a 99% selectivity in relation to methanol, and 90% in relation to CO, not all the reactants eventually
combine to produce acetic acid. Other important side reactions occurring in the system generate acetic
acid, methyl acetate and propionic acid.
4.1.1.1 Propionic acid
𝐶2 𝐻6 𝑂 + 𝐶𝑂 → 𝐶3 𝐻6 𝑂2
Propionic acid is the major liquid by-product within the system. Ethanol impurity is present in the methanol
streams that reacts with some of the unreacted carbon monoxide. The reaction is kinetic and its value have
been researched in literature.
4.1.1.2 Water – Gas shift
𝐻20 + 𝐶𝑂 → 𝐶𝑂2 + 𝐻2
The reaction once again involves part of the unreacted carbon monoxide reacts with water to generate
hydrogen and carbon dioxide. This reaction is equilibrium limited.
4.1.1.3 Methyl acetate
𝐶𝐻3 𝑂𝐻 + 𝐶𝐻3 𝐶𝑂𝑂𝐻 → 𝐶𝐻3 𝐶𝑂𝑂𝐶𝐻3 + 𝐻2 𝑂
A fraction of the 1% unreacted methanol combines with the product (acetic acid) to generate methyl acetate
and water. This reaction is equilibrium limited.
4.1.1.4 Methyl iodide
𝐶𝐻3 𝑂𝐻 + 𝐻𝐼 → 𝐶𝐻3 𝐶𝑂𝑂𝐶𝐻3 + 𝐻2 𝑂
A fraction of the 1% unreacted methanol combines with the hydrogen iodide to generate methyl iodide and
water. This reaction is equilibrium based.
An Aspen PLUS simulation has been modelled based on:
• Information researched in literature
• Appropriate calculations

17
Methanol feed
6' SS
P-102
R101
PI
103
FI
103
FCV-102
FCV-101
101 101
FC FT
102 102
PCV-103
P-103
Methanol storage tank
P-101
FC FT
P-201
LCV-201
5.5" SS
P-202
103
FI
103103
PCPT
203
FI
202
LIC
201201
FTFC
PRV-201
201
PI
102 102
TI PI
103
TI
PRV-201
203
PI
FCV-205
205
205
FT
FC
H2O supply
6" SS
LAH
LAL
204 204
TI PITLH
TLL
PLH
PLL
PLH
PLH
Off to flash tank
From flash tank
Off-to scrubber
From drying column
3/8" - SS
3" - SS
80" - SS
6" – Hastelloy
6" – Hastelloy
12" – CS
CO
Supply
Figure [4.1]: P&ID of Reactor.
R-201
4.1.2 Reactor P&ID

18
4.1.3 Design Method
4.1.3.1 Reactants and product
Since the reactants are in different phases, liquid and gas respectively, it is advantageous to insert the gas
through a sparger to facilitate even diffusion throughout the liquid. The sparger should be designed
separately.
Stream Density (kg/m3
)
Methanol 791
CO 1.14
Acetic Acid 1005
Table [4.2] – Composition of stream and respective densities.
Table [4.3] – Material balance for reactor.
Substream:MIXED CO Methanol Water CSTR-LIQ CSTR-VAP WATER-RCY SCRUB-RCY
MassFlowkg/hr
Methanol 0.00 27044.00 0.00 56.15 1.15 0.00 57.29
CO 25209.36 0.00 0.00 207.44 200.88 0.00 105.97
AceticAcid 0.00 0.00 0.00 98543.49 611.25 29094.25 9045.00
Water 0.00 0.00 720.40 234.38 3.19 0.00 214.13
Ethanol 0.00 5.93 0.00 0.00 0.00 0.00 0.00
PropionicAcid 0.00 0.00 0.00 487.99 2.02 0.00 2.02
CO2 0.00 0.00 0.00 6256.54 947.56 0.00 4773.63
H2 0.00 0.00 0.00 8.69 72.79 0.00 1.60
Methylacetate 0.00 0.00 0.00 1243.25 24.77 0.00 1268.02
Hydrogenchloride 0.00 0.00 0.00 0.00 0.00 0.00 0.00
Methyliodide 0.00 0.00 0.00 16.19 0.43 0.00 13.38
TotalFlowmol/hr 900000.00 850000.00 40000.00 1849262.00 75922.76 484500.60 307898.00
TotalFlowkg/hr 25209.36 27319.02 720.61 109343.00 1924.24 29095.50 17368.82
volumetricflowl/hr 22299120.00 34446.27 724.99 126869.52 89522.94 419083.02 17696.84
TemperatureC 25.00 25.00 25.00 160.00 160.00 220.04 69.71
Pressurebar 1.00 1.00 1.00 30.40 30.40 32.00 59.00
VaporFrac 1.00 0.00 0.00 0.00 1.00 1.00 0.00
LiquidFrac 0.00 1.00 1.00 1.00 0.00 0.00 1.00
SolidFrac 0.00 0.00 0.00 0.00 0.00 0.00 0.00
Concentration(mol/l) 25.00 25.00 1.00 14.58 0.85 1.16 17.40
Densitygm/cc 1.13 793.09 993.96 861.85 21.49 69.43 17.40
INLET OUTLET RECYCLE

19
4.1.3.2 Variations
The material and energy balances produced refer to steady state operating conditions. There are some
situations, however, in which the system does not operate under steady state conditions, including:
• Start-up and shutdown
• Filling and emptying
• Isolation
• Preventative or corrective maintenance
• Failure or loss of process variation
• Extreme ambient conditions (temperature, severities, pressures)
4.1.3.3 Energy Balance and Heat of reaction
As acknowledged from literature, the methanol carbonylation reaction is exothermic. The heat of reaction is
therefore calculated with the aid of enthalpies of each stream (reactants and products) as seen by Figure
4.2.
Figure [4.2]: Energy balance of inlet, outlet and recycle stream.
The negative sign of heat of reaction means that the system releases energy in the form of heat. This heat
of reaction has been calculated within the simulation through the basic formula:
△ 𝐻 𝑃𝑟𝑜𝑑𝑢𝑐𝑡𝑠 − △ 𝐻 𝑅𝑒𝑎𝑐𝑡𝑎𝑛𝑡𝑠 = 𝑞 = 𝐸𝑛𝑒𝑟𝑔𝑦 𝑔𝑒𝑛𝑒𝑟𝑎𝑡𝑒𝑑 𝑖𝑛
𝑐𝑎𝑙
𝑚𝑜𝑙
𝑜𝑓 𝑝𝑟𝑜𝑑𝑢𝑐𝑡
The specific enthalpies of each stream have been obtained by the simulation. As far as simulation is
concerned, the water inlet stream’s purpose is to enable the catalyst and promoter activity. It is not involved
in any chemical reaction and it is in fact separated later on in the process from the product stream, and
recycled back into the reactor continuously. For this reason it is possible to state that its energy contribution
is negligible and not relevant to take into account for energy balance purposes. The same principle is
applied to the recycling streams connecting the reactor to scrubber and reactor to drying column
respectively, therefore they have not been included in this energy balance.
For the purpose of design optimization and practice, heat needs to be continuously removed from the
system and applied to an appropriate heat integration system.
CO Methanol Water CSTR-LIQ CSTR-VAP
Enthalpy (cal/mol) -26401.54 -57103.93 -68275.29 -108490.00 -39291.50
Flowrate (mol/hr) 900000 850000 40000 1849262 75922.76
Total enthalpy (cal/hr) -23761386000 -4.8538E+10 -2.73E+09 -2.00626E+11 -2983119125
Total enthalpy (cal/hr)
Enthalpy of reaction (cal/hr)
Enthalpy of reaction (kJ/hr)
-72299726500 -2.0361E+11
-1.3131E+11
-5.49E+11

20
4.1.3.4 Choice of reactor
The nature of the reaction involved in the process necessitates agitation within the system. The exothermic
nature of the main reaction emphasizes that a vessel with good temperature control is desirable.
Furthermore, the fact that a continuous operation of the vessel is required and a simple adaptation to two-
phase reaction is possible implies that a CSTR is the most suitable choice of reactor in the Monsanto
methanol carbonylation. Advantages of CSTR’s are:
• Constant operation in continuous system.
• High degree of temperature and process control.
• Simplicity of construction.
• Easy adaptation to two phase reactions (liquid-gas reaction).
• Easy maintenance / clean-up operations.
Assumptions that will be useful in the later stage of the reactor design can be made use on the CSTR of the
process:
• Steady state conditions with constant inlet (reactants) and outlet flow (products).
• Uniform stream composition inside and outside the reactor.
• Complete and uniform mixing.
4.1.4 Reactor Specification
Reactor – Design Data
Vessel volume 138.42 m3
Vessel shell diameter (internal) 4.45 m
Internal pressure 30 bara = 30.59 kg/cm2
External pressure 1 bara
Design pressure (10% of Operating pressure) 32.12 kg/cm2
= 33 bara = 3.3 N/mm2
Allowable stress (Hastelloy B) 351.63 N/mm2
Hydrostatic test pressure 39.77 kg/cm2
Density of material 9022 kg/cm3
Corrosion allowance (CA) 4 mm
A thorough calculation of reactor design can be seen in Appendix B.

21
4.1.4.1 Choice of material
Hastelloy B-2 – (65% Ni, 28& Mo, 5% Fe) required to hold resistance against the corrosion of hydrogen
iodide and acid (Sinnott et al., 2009, p986). Its physical and chemical properties appear to be specifically
suitable as a choice for the CSTR of this system.
𝐻𝑎𝑠𝑡𝑒𝑙𝑙𝑜𝑦 𝑌𝑖𝑒𝑙𝑑 𝑆𝑡𝑟𝑒𝑛𝑔𝑡ℎ = 51 𝑘𝑠𝑖 = 2179 𝑘𝑔𝑓/𝑚2
(Alloys and Producer, 2014)
4.1.4.2 Vessel support
A support skirt will be required for the reactor. The design of such implementation will have to be performed
separately and in accordance to the vessel’s specifications.
4.1.4.3 Piping sizing
Pipeline Flowrate
(m3
/s)
Velocity
(m/s)
Cross -
sectional
area (m2
)
Diameter (mm) Equivalent NPS (in)
CO feed to
reactor
6.1942 2.000 3.0971 1986.3 80
Methanol feed
to reactor
0.0096 2.000 0.0048 78.1 3
Water feed to
reactor
0.0002 2.000 0.0001 11.3 0.375
Reactor to
flash tank
0.0304 2.000 0.0152 139.1 6
Reactor to
scrubber
0.0249 2.000 0.0124 125.9 6
Drying column
to reactor
0.1164 2.000 0.0582 272.3 12
Scrubber to
reactor
0.0049 2.000 0.0025 2.5 12
Table [4.4]: Representation of pipeline diameter upon data obtained from Aspen Plus simulation and conversion
(mm to NPS) (Perry et al., 1997).

22
The feed of CO to the reactor pipeline represents an offset value, which is due to the high volumetric
flowrate as a result of its gas-phase nature, leading to very low density. On the basis that a constant
velocity of 2 m/s is assumed, the CO feed pipeline will be therefore be significantly larger.
4.1.4.4 Nozzles
Divergent nozzles given a size margin of 15% and distance between nozzle entrance and flange distance
assumed to be twice of the nozzle’s diameter. Dimensions of nozzle to appropriate pipeline is given in
Table 4.5.
Table [4.5]: Representation of nozzles in relation to pipeline upon data obtained from Aspen Plus simulation and
conversion (mm to NPS) (Perry et al., 1997).
4.1.4.5 Heat dissipation and vessel insulation
Similarly to most systems to which the laws of thermodynamics can be applied to, there is dissipation of
energy in the form of heat to the surrounding areas. This rate of energy exchange is governed by Fourier’s
Law, which states:
𝑄 = ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟 = 𝑘 𝐴
dT
𝑑𝑠
Where:
𝑄 = ℎ𝑒𝑎𝑡 𝑡𝑟𝑎𝑛𝑠𝑓𝑒𝑟
k = thermal conductivity of material
A = area of the vessel
𝑑𝑇 = temperature gradient
Pipeline Nozzle Diameter (mm) Equivalent NPS (in) Nozzle entrance to flange
(m)
CO feed to
reactor
A 2284.23 88 4.57
Methanol feed to
reactor
B 89.78 5 0.18
Water feed to
reactor
C 12.99 0.375 0.03
Reactor to flash
tank
D 159.95 8 0.32
Reactor to
scrubber
E 144.73 6 0.29
Drying column to
reactor
F 313.15 14 0.63
Scrubber to
reactor
G 64.35 2.5 0.13

23
𝑑𝑠 = thickness of material
𝑄 = 67,301,345 kJ/hr (𝑆𝑒𝑒 𝐶𝑎𝑙𝑐𝑢𝑙𝑎𝑡𝑖𝑜𝑛𝑠 𝑖𝑛 𝐴𝑝𝑝𝑒𝑛𝑑𝑖𝑥 [𝐵])
Normally the issue of heat dissipation from a vessel can be overcome by applying an appropriate insulation
system on the internal fitting of the vessel; in this case, however, the heat dissipated is negligible in relation
to the heat of reaction produced within the reactor:
𝐻𝑒𝑎𝑡 𝑜𝑓 𝑟𝑒𝑎𝑐𝑡𝑖𝑜𝑛 𝑝𝑟𝑜𝑑𝑢𝑐𝑒𝑑 = 5.49 × 1011
kJ/hr
𝑄 = ℎ𝑒𝑎𝑡 𝑑𝑖𝑠𝑠𝑖𝑝𝑎𝑡𝑖𝑜𝑛 = 67,301,345 = 6.7 × 107
kJ/hr
Therefore by taking into account this specific system, it is fair to assume that heat dissipation is not relevant
in relation to the rate of heat of reaction produced, so the conclusion is that no insulation is required for this
vessel.
4.1.4.6 Shutdown
Shutdown procedures are required whenever an emergency shutdown or maintenance occurs. The
standard safety protocols must be followed, where the first step is to decrease production rate at constant
intervals, thus allowing the progressive reduction of pressure and temperature. When conversion reaches
0, the reactor will need to get emptied. In order to do so, the vessel is continuously purged with an inert
substance, nitrogen, to prevent formation of oxygen or other reactions for when the shell comes into
contact with the atmosphere.
4.1.4.7 Process safety
The choice of material, Hastelloy, is the first application of safety factors within the design. As a result of the
highly corrosive nature of the iodides present, hydrogen and methyl iodide, Hastelloy represents the most
suitable material choice. A high corrosion allowance (4mm) has thus been implemented. Parameters such
as temperature and pressure, other than level of fluid within the vessel are directly related to the safety
procedures, therefore appropriate process control measures have been employed to ensure a high level of
process safety in the reactor.
4.1.4.8 Process control
Methanol feedlines have been integrated with temperature and pressure indicators alongside a flow control
valve with an appropriate transmitter due to flowrate being the most relevant parameter to control in order
to allow the reaction to occur currently. A backup pump has also been employed in the methanol feed
pipelines, in case of failure of the first one. Moreover, backup flow control valves have been included in
both the CO and methanol feeds.

24
Pressure control and a transmitter have been included in the reactor to keep a constant monitoring of the
pressure within the vessel, thus overpressure or under pressure is prevented which would affect the rate of
reaction (30 bar). A level indicator controller, with high and low alarms, have been selected to monitor the
reactor’s fluid level; they control the flow valve located in the pipeline between the liquid phase of the
reactor and the flash tank in order to maintain fluid level within vessel specification, a maximum 70% of
vessel’s volume is recommended).
Isolation valves have been fitted around the vessel to ensure a high level of control in relation to safety and
failures of the system, allowing quick shutdown of the unit operation. The pressure reduction valve located
in the pipeline which connects the liquid stream of the rector to the flash tank is fitted in order to flash the
content and reduce pressure from 30 bar to 1 bar. The other streams feeding from and to the reactor have
been fitted with appropriate flow controllers and transmitters to monitor the overall flow within the system
4.1.5 Agitator Specification
Agitators are required to increase the transfer of material within the vessel and create uniform temperature
within the reactor.
In order to evaluate which type of agitator is most suitable for the reaction, it is possible to observe a
correlation between viscosity, volume of the tank and type of agitator.
Figure [4.3]: Aspen Plus simulation data.
Stream Flowrate (kg/hr) Mass fraction
avg. viscosity (N-s/m^2)
avg.viscosity (N-s/m^2)( fraction)
CO 25209.36 0.119486656 1.77E-05 2.11E-06
Methanol 27319.02 0.129485966 0.000540605 7.00E-05
Water 720.6112 3.42E-03 0.000912531 3.12E-06
CSTR-LIQ 109343 5.18E-01 0.000132853 6.89E-05
CSTR-VAP 1924.236 9.12E-03 1.83E-05 1.67E-07
WATER-RCY 29095.5 0.137906079 1.57E-05 2.16E-06
SCRUB-RCY 17368.82 0.082324272 0.000119604 9.85E-06
Total 210980.5472 1 1.56E-04
0.156 mpa.s
LIQUID VISCOSITY 0.000793062

25
Figure [4.4]: Aspen Plus simulation data.
The values for viscosity have been collected from the simulation on Aspen PLUS. Through implementation
of flowrates for all the streams connected to/from the reactor, a mass flowrate has been calculated, hence
leading to calculate the total liquid viscosity of the mixture, and correlating Figure 4.3, this leads to:
𝑉𝑜𝑙𝑢𝑚𝑒 𝑜𝑓 𝑣𝑒𝑠𝑠𝑒𝑙 = 138 𝑚3
𝐿𝑖𝑞𝑢𝑖𝑑 𝑣𝑖𝑠𝑐𝑜𝑠𝑖𝑡𝑦 = 7.9 × 10−3
𝑁𝑠/𝑚2
𝐺𝑟𝑎𝑝ℎ 𝑖𝑛𝑡𝑒𝑟𝑐𝑒𝑝𝑡𝑖𝑜𝑛 = 𝑃𝑟𝑜𝑝𝑒𝑙𝑙𝑒𝑟 𝑜𝑟 𝑡𝑢𝑟𝑏𝑖𝑛𝑒 (𝑆𝑒𝑒 𝐶𝑎𝑙𝑐𝑢𝑙𝑎𝑡𝑖𝑜𝑛𝑠 𝑖𝑛 𝐴𝑝𝑝𝑒𝑛𝑑𝑖𝑥 𝐵)
Figure [4.5]: Impeller type for different mixture viscosities (Sinnott et al., 2009).
By analysing Figure 4.5, it is possible to evaluate that flat-blade turbine is the most suitable one in relation
to the liquid viscosity of 7.9 × 10-3
Ns/m2
.
Stream Flowrate (kg/hr) Mass fraction
avg. density (kg/m3)
avg.density (kg/m3)( fraction)
CO 25209.36 0.119486656 1.13051 1.35E-01
Methanol 27319.02 0.129485966 793.0907 1.03E+02
Water 720.6112 3.42E-03 993.957 3.39E+00
CSTR-LIQ 109343 5.18E-01 861.8497 4.47E+02
CSTR-VAP 1924.236 9.12E-03 21.4943 1.96E-01
WATER-RCY 29095.5 0.137906079 69.4265 9.57E+00
SCRUB-RCY 17368.82 0.082324272 17.3984 1.43E+00
Total 210980.5472 1 5.64E+02
564 kg/m3

26
Using ratios given in literature allows to specify the dimensions, location and characteristics of impellers
and blades have been identified. In order to guarantee a good level of reaction control in relation to heat
within the vessel, baffles have been integrated to the design.
Due to the fact that µ < 500 mPa.s, the type of agitator should be either a propeller or a turbine. It is
assumed that the agitation required is mild for such homogenous reaction. Therefore, assumed rotational
speed falls within the ‘low’ category, 200 rpm is the estimated rotational velocity required (Carpenter, 2010).
Agitator – Design Data
Diameter 1.48 m
Height above vessel bottom 1.48 m
Blade length 0.37 m
Blade width 0.45 m
Baffled width 0.45 m
Baffled height 8.9m
Liquid depth 4.45 m
Number of blades 6
Number of baffles 4
Diameter of shaft 0.053 m
Power required 194 hp
A thorough calculation of agitator design can be seen in Appendix B.
4.1.6 Conclusion
The parameters, data and information researched about the process have been defined and simulated on
appropriate software (Aspen Plus). This allowed to facilitate the collection of results in relation to mass and
energy balances in order to define the main parameters of the vessel, allowing a 20% overdesign for
potential expansion. A detailed mechanical design of the reactor including its minor components has been
completed, defining the dimensions of turbine, blades, shaft and the power required by the agitator. In
addition, a comprehensive stress analysis has outlined that the reactor’s mechanical design is suitable for
industrial application. Time taken to reach steady state conditions as well as an analysis of shutdown
operations, process control and safety has been implemented throughout the design. Furthermore, a cost

27
estimation has been evaluated in relation to the plant location, time and local currency. The final cost for
the unit is approximately £800,000 – excluding delivery and installation costs.

28
4.1.7 Engineering Drawing of Reactor
Figure [4.6]: Mechanical design of reactor

29
4.2 Flash Tank
4.2.1 Introduction
The flash drum is a vapour liquid separator, its role is to split the mixture of the vapour-liquid mixture fed
from the reactor. The vapour stream is released from the top of the drum. The liquid stream leaves through
the bottom of the drum containing the Rhodium catalyst which is then recycled back to the reactor to be
reused to aid reactions within the reactor. The design approach was to specify the flash tanks operating
conditions and physical attributes (addition of demister and diameter and tank length) that directly affect the
cost of equipment and operating costs. The orientation of the vessel will be vertical as its ideal for high flow
rates, the vertical separator the process is more economical compared to the horizontal separator
(Monnery and Svreck, 1993)
A demister pad is a device with metal mesh like structure that eliminates the possibility of liquid entrainment
within a pressure vessel. Entrainment is the entrapment of one phase within another, within the flash drum
liquid droplets can be entrained within vapour and leave liquid droplets within the vapour stream. To
prevent liquid entrainment, the velocity of vapour stream must be kept low to allow the water droplets to
disengage for the vapour stream and drop back down to the liquid pool at the base of the vessel. If the
operation requires a high vapour velocity the demister pad acts as an effective entrainment separator
(Basic et al., 2013).As the vapour travels through the mesh wiring pad the stream lines are deflected,
however the kinetic energy of the liquid droplet entrained within the vapour are too high to follow the
streamline, they become impinged in the wires. The liquid droplets then coalesce forming a liquid layer on
the surface of the wires. The droplets then detach from the pad. Due to the orientation of the vessel
(vertical) the liquid droplets will be captured and be drained back and form large droplets that can drop
from the upstream face of the wire mesh pad ( Al-Deffeeri et al., 2000).
Demister pads increase the efficiency of vapour liquid separation efficiency. Flash drums that use gravity
separation (without the demister) are dependent on a high residence time to separate the liquid from the
vapour. The more time needed for the mixture to separate, the higher the energy cost to run the flash drum
thus the plant throughput will be lower per day hence reducing revenue per day. The demister pad allows
the same degree of separation to be carried out in a smaller vessel, the reduction of volume reduces the
weight of the vessel which directly minimizes the cost of the vessel shell (Sinnott et al., 2009)).The internal
diameter is dependant of the vessel is dependent on the terminal velocity of the particles

30
4.2.2 Flash Tank P&ID
F-201
V-206
FIC
LI
V-211
FI FIC
R-201
V-210
V-209
V-208
V-207
FI
FI FIC
FIC
R-201
DC-301
Figure [4.7]: P&ID of Flash Tank

31
4.2.3 Design Method
This specific sizing methodology is adopted from “two phase separators within the right limits” published in
the “Chemical engineering progress synopsis series” (1993) the calculation initiates by the finding the
diameter of the vessel. In order to do so, vertical terminal vapour velocity, QT, is determined by obtaining
the K value using Table C1 (See Appendix C). Subsequently, QV, vapour volumetric flow rate is calculated.
The internal vessel diameter, DVD, is estimated, whilst adding 6 inches to the figure obtained to
accommodate the support for the mist eliminator. Referring to Table C2, hold up time and surge volume
relative to a “Feed to column separator” are selected. Further referring to Table C3, low liquid level height,
HLLL, is obtained, thus distance from the low liquid level, HLLL to the normal liquid level, HNLL, is estimated.
This value must be minimum of 1ft. Consecutively, the height between normal liquid level, HNLL to high liquid
level HHLL, must be 6 inches minimum. Henceforth, the height from high liquid level to the centre line of inlet
nozzle is estimated.
The disengagement height from the centre line of inlet nozzle to the bottom of demister pad is then
determined and assumption of the height of the mist eliminator pad, HME, is 6 inches and 1ft is taken from
the top of the mist eliminator to the tangent line of the flash drum.

32
Mole Flow kmol/hr Flash in Flash vapour stream Flash liquid stream
METHANOL 1.752356 1.752356 0
CARBON MONOXIDE 7.405816 7.405816 0
ACETIC ACID 1641.024 1476.921 164.1024
WATER 13.01377 11.71239 1.301377
ETHANOL 1.35E-05 1.22E-05 1.35E-06
PROPIONIC ACID 6.588246 5.929421 0.6588246
CARBON DIOXIDE 142.194 142.194 0
HYDROGEN 4.304404 4.304404 0
METHYL ACETATE 16.78482 16.78482 0
HYDROFEN CHLORIDE 7.75E-08 7.75E-08 0
METHY-IODIDE 16.19486 12.95588 3.238971
Total Flow kmol/hr 1849.262 1679.96 169.3015
Total Flow kg/hr 109343 98955.71 10386.75
Total Flow l/min 266948 264110 176.0507
Temperature C 104.4454 104.4454 104.4454
Pressure bar 1 1 1
Vapor Frac 0.3835523 0.417973 0
Liquid Frac 0.6164477 0.582027 1
Solid Frac 0 0 0
Enthalpy cal/mol -107770 -107520 -110540
Enthalpy cal/gm -1822.75 -1825.307 -1801.79
Enthalpy cal/sec -55362000 -50173000 -5198500
Entropy cal/mol-K -57.37749 -56.42035 -67.50691
Entropy cal/gm-K -0.9704007 -0.9578422 -1.100346
Density mol/cc 0.000115457 0.000106014 0.0160277
Density gm/cc 0.00682671 0.00624461 0.983311
Average MW 59.12762 58.9036 61.35062
Liq Vol 60F l/min 1766.089 1603.876 162.2127
Table [4.6]: Material Balance for Flash Tank.
4.2.4 Flash Tank Specification
Flash Tank – Design Data
Design pressure 1.1 bar
Design temperature 114.89 o
C
Pressure 1 bar
Temperature 104.445 o
C
Vapour volumetric flow rate 4.4 m3
/min
Liquid volumetric flow rate 1.68 m3
/min
Vapour density 6.24 kg/m3
Liquid density 983.31 kg/m3

33
The chosen material of construction is Hastelloy-B-3 (65% Ni, 28% Mo, 5% Fe) due to the corrosive nature
of Hydrogen Iodide and acid. The minimum allowable diameter of the vessel has to be large enough to slow
down the gas below the velocity which the particles will settle out (Sinnott et al., 2009).
Following calculations from (Monnery and Svreck, 1993):
𝑇𝑒𝑟𝑚𝑖𝑛𝑎𝑙 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 𝑈 𝑇 = 1.32 𝑚/𝑠
𝑇𝑒𝑟𝑚𝑖𝑛𝑎𝑙 𝑣𝑎𝑝𝑜𝑢𝑟 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 𝑈 𝑉 = 0.992 𝑚/𝑠
𝑉𝑎𝑝𝑜𝑢𝑟 𝑣𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 = 𝑄 𝑉 = 4.4 𝑚3
/𝑠
The flash drum has a mist eliminator, therefore 6 inches are added to accommodate a support ring and
rounding up to the next 6-inch increment to obtain the external diameter. Thus:
𝐼𝑛𝑡𝑒𝑟𝑛𝑎𝑙 𝑣𝑒𝑠𝑠𝑒𝑙 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 𝐷 𝑉𝐷 = 2.53 𝑚
𝐸𝑥𝑡𝑒𝑟𝑛𝑎𝑙 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 𝐷 = 3.66 𝑚
Hold up time is the time it takes for the normal liquid level to reach the lower liquid level to empty whilst
keeping a normal outlet flow rate with no feed entering the vessel. Thus:
𝐻𝑜𝑙𝑑 𝑢𝑝 𝑡𝑖𝑚𝑒 = 𝑇 𝐻 = 5 𝑚𝑖𝑛𝑢𝑡𝑒𝑠
𝐻𝑜𝑙𝑑 𝑢𝑝 𝑣𝑜𝑙𝑢𝑚𝑒 = 𝑉 𝐻 = 8.39 𝑚3
The surge time is the time it takes for the normal liquid level to rise from normal liquid level to maximum
when keeping normal feed flow rate and no outlet flow. Therefore:
𝑆𝑢𝑟𝑔𝑒 𝑡𝑖𝑚𝑒 = 𝑇𝑆 = 3 𝑚𝑖𝑛𝑢𝑡𝑒𝑠
The volume of liquid between the highest liquid level and the normal liquid level
𝑆𝑢𝑟𝑔𝑒 𝑣𝑜𝑙𝑢𝑚𝑒 = 𝑉𝑆 = 5.032 𝑚3
The total height of the vertical flash drum is the sum of the HLLL +HH +HS + HLIN +HD +HME + 1ft. The height of
the vessel outlet at the top of the vessel must be sufficient so the liquid droplets can disengagement from
the vapour. Henceforth, liquid heights within vessel:
𝐿𝑜𝑤 𝑙𝑖𝑞𝑢𝑖𝑑 𝑙𝑒𝑣𝑒𝑙 ℎ𝑒𝑖𝑔ℎ𝑡 = 𝐻𝐿𝐿𝐿 = 0.1524 𝑚
𝐻𝑒𝑖𝑔ℎ𝑡 𝑏𝑒𝑡𝑤𝑒𝑒𝑛 𝐻𝐿𝐿𝐿 𝑎𝑛𝑑 𝐻 𝑁𝐿𝐿 = 𝐻 𝐻 = 1.667 𝑚
𝐻𝑒𝑖𝑔ℎ𝑡 𝑏𝑒𝑡𝑤𝑒𝑒𝑛 𝐻 𝑁𝐿𝐿 𝑎𝑛𝑑 𝐻 𝐻𝐿𝐿 = 𝐻𝑆 = 1 𝑚
𝐻𝑒𝑖𝑔ℎ𝑡 𝑓𝑟𝑜𝑚 𝐻𝐿𝐿𝐿 𝑡𝑜 𝑡ℎ𝑒 𝑐𝑒𝑛𝑡𝑟𝑒 𝑙𝑖𝑛𝑒 𝑜𝑓 𝑖𝑛𝑙𝑒𝑡 𝑛𝑜𝑧𝑧𝑙𝑒 = 𝐻 𝐷 = 3.66 𝑚
𝐻𝑒𝑖𝑔ℎ𝑡 𝑓𝑟𝑜𝑚 𝐻 𝐻𝐿𝐿 𝑡𝑜 𝑐𝑒𝑛𝑡𝑟𝑒 𝑙𝑖𝑛𝑒 𝑜𝑓 𝑛𝑜𝑧𝑧𝑙𝑒 = 𝐻𝐿𝐼𝑁 = 4.062𝑚

34
𝑀𝑖𝑠𝑡 𝑒𝑙𝑖𝑚𝑖𝑛𝑎𝑡𝑜𝑟 ℎ𝑒𝑖𝑔ℎ𝑡 = 𝐻 𝑀𝐸 = 0.1524 𝑚
𝐻𝑒𝑖𝑔ℎ𝑡 𝑏𝑒𝑡𝑤𝑒𝑒𝑛 𝑚𝑖𝑠𝑡 𝑒𝑙𝑖𝑚𝑖𝑛𝑎𝑡𝑜𝑟 𝑎𝑛𝑑 𝑡𝑜𝑝 𝑡𝑎𝑛𝑔𝑒𝑛𝑡 𝑙𝑖𝑛𝑒 𝑜𝑓 𝑣𝑒𝑠𝑠𝑒𝑙 = 𝐻 𝑇𝑇 = 0.305 𝑚
𝑇𝑜𝑡𝑎𝑙 𝑓𝑙𝑎𝑠ℎ 𝑑𝑟𝑢𝑚 ℎ𝑒𝑖𝑔ℎ𝑡 = 𝐻𝐿𝐿𝐿 + 𝐻 𝐻 + 𝐻𝑆 + 𝐻𝐿𝐼𝑁 + 𝐻 𝐷 + 𝐻 𝑀𝐸 + 𝐻 𝑇𝑇 = 10 𝑚
When calculating the corrosion wall thickness of a vessel the corrosion allowance must be taken into
consideration. The corrosion allowance is the amount of Hastelloy material available for corrosion without
disturbing the amount of pressure the vessel can contain (Sinnott et al., 2009). Thus:
𝐶𝑜𝑟𝑟𝑜𝑠𝑖𝑜𝑛 𝑎𝑙𝑙𝑜𝑤𝑎𝑛𝑐𝑒 = 4 𝑚𝑚
Leads to:
𝑊𝑎𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = 3.6 𝑚𝑚 + 4 𝑚𝑚
= 7.6 𝑚𝑚
𝑇ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 𝑜𝑓 𝑣𝑒𝑠𝑠𝑒𝑙 ℎ𝑒𝑎𝑑𝑠 = 7.6 𝑚𝑚 + 4 𝑚𝑚
= 11.49 𝑚𝑚
The internal diameters of the Hastelloy pipes leaving and entering the flash drum were calculated using
Sinnott and Towler, the method is further discussed in Appendix [C].
Pipeline Flowrate (m3
/s) Velocity (m /s) Cross
sectional area
(m2
)
Required diameter (m)
Reactor to Flash
drum
4.45 2.00 2.225 1.68
Flash drum to light
ends column
4.40 2.00 2.2 1.67
Flash drum to
reactor
0.0029 2.00 1.45 x 10-3 0.043
Table [4.7]: Pipeline sizing for Flash Tank
4.2.4.1 Process controls and safety
On the left hand side of the flash drum there is a level indicator which will set off an alarm when liquid level is below a
certain point, this is to avoid pump damage. When the pump is pumping air and not fluid it can lead to cavitation and
produce loud noises. The pressure indicator send sends message to the alarm system if the pressure within the
vessel is over the maximum pressure of 1.1 bar. The pressure within the flash tank is relatively low compared to the
reactor and it’s very rare that it tends to overheat because the heat exchanger reduces the temperature before all
products reach the flash drum.
Going beyond this pressure can result in boiling of the liquid within the vessel and increase of temperature. The
contents within the vessel are highly hazardous and flammable the build-up of pressure can cause an explosion and
put all site workers at risk but this is highly unlikely. If the flash tank temperature did increase it would be due to a
faulty heat exchanger feeding through high temperature streams, so to reduce the temperature the reactor must be
cooled down. So a pipeline must be installed and redirected, while the faulty heat exchanger is fixed
4.2.5 Conclusion

35
The flash drum unit was designed with 20 % overdesign ass specified in the process specification. The
P&ID provided gives detail to the indicators and control systems essential for process safety and efficient
production. The Engineering drawing is cross sectional representation of the essential internal units and
recommended liquid levels within the flash drum. The final cost for the unit is approximately £67,000.

36
4.2.6 Engineering Drawing of Flash Tank
Fig [4.8]: Mechanical design of Flash Tank.

37
4.3 Drying column
4.3.1 Introduction
This section of the design project report provides a detailed design description of the drying column in the
process. The process specification requires a yield of 400,000 tonnes per year and 99% purity of acetic
acid, therefore it is mandatory that the drying column is optimized to meet the client’s specification. The
objective of this section is to calculate the operating condition and physical parameters required to optimise
the process in order to meet the client specification; for example the column diameter, height, thickness,
ends and tray sizing. The column parameters are shown in Table 4.6.
Feed rate, F kg/hr 83213.17
Feed Composition
Acetic acid 99.6 kmol%
Water 3.56 x10-11 kmol %
Propanoic acid 0.428 kmol%
Feed Temperature, oC 118
Column operating pressure, bar 1
Column reboiler Partial Reboiler
Column condenser Partial Condenser
Distillate composition, XD 0.992
Bottoms composition, XB 0.038
Table [4.7]: Specified column parameters.

38
4.3.2 Drying Distillation Column P&ID
Figure [4.9]: P&ID of Drying Column
DC-401

39
4.3.3 Design Method
Kinetic and thermodynamic data were collected from literature for water and acetic acid whilst obtaining
stream properties from the simulation on Aspen Plus, thus allowing to identify XD and XB in the vapour and
liquid streams. The subsequent data allows the determination of the number of trays in the column by
utilising the McCabe – Thiele Method for binary mixtures and thus feed position and reflux conditions are
estimated. In order to propose a viable implementation of a mechanical design, dimensions of the column
need to be determined, for example diameter and height, as well as selecting a suitable materials of
construction, a preliminary mechanical design of the drying column, which comprises of column design,
plate design and general arrangements and finally estimate a proposed cost of the column, including
capital and operating cost (Sinnot et al., 2009).
4.3.4 Drying Distillation Column Specification
The final drying distillation column specification is based on the calculations in Appendix D. The drying
distillation column is represented diagrammatically in Figure 4.8.
The purpose of the drying column in this acetic acid synthesis process is to increase the acetic acid purity
via separating methanol, ethanol, methyl acetate and water from the product stream and recycle these
undesired compounds back in to the reactor, consequently increasing the purity of the product stream. A
sketch of the column and plate is shown in Figure 4.8.
Drying Column – Design Data
Working Pressure 1.1 bar
Inside Diameter (Di) 1850 mm
Material of construction 305 Stainless Steel
Allowable Stress 515 N/mm2
Density of material 8027.172 kg/m3
Design pressure 1.1 bar
Height of column 23.85 m
Area of column 2612.81 m2
Thickness of column 12 mm
End selection Torispherical
End thickness 20 mm
Number of trays 34

40
Feed entry 16
Plate spacing 0.7 mm
Hole pitch (rectangular) diameter 5 mm
Tray thickness 3.5 mm
Packing size 75 mm
Pipe diameter Feed 216 mm
Top 178 mm
Bottom 148 mm
The design considerations were made based on the specification provided from the client, the
following should be noted in the design.
Drying Distillation Column – Design Considerations
Cost of shell and trays £364,000
Cost of reboiler £22,000.00
Cost of condenser £16,500.00
Dead-weight of shell 122 kN
Weight of plates 109.67 kN
Weight of insulation 5.6 kN
Weight of vessel 237.27 kN
Wind loading 31.94 N/mm2
Bending moment 57192.3 N/mm2
Longitudinal stresses 84.8 N/mm2
Circumferential stresses 48.4 N/mm2
4.3.5 Conclusion
The following design specification on this unit complies with the necessary design intent specification. In
addition, the specification fulfils the plant debottleneck allowances of a 20% overdesign.

41
4.3.6 Engineering Drawing of Drying Distillation Column
Figure [4.10]: Mechanical Design of Drying Column

42
4.4 Heavy – Ends Distillation column
4.4.1 Introduction
This column is designed to separate the unwanted propionic acid produced as part of the process from the
desired acetic acid produced, to obtain a purity of greater than 99.9% which is required by the design brief
at a capacity of 400,000 tonnes per year.
Temperature = 125 o
C (398K)
Operating conditions Pressure = 1 atm (101325 Pa)
Reflux ratio = 17
For the purposes of the calculations, the average internal temperature of the column was assumed to be 125
o
C (398K) to ensure that the majority of the acetic acid, and only a minimal amount of propionic acid was in
the vapour phase in addition to operating at 1 atm (101325 Pa). The reflux ratio of 17 was taken from the
simulation produced as part of this project on Aspen Plus, as this gave the desired quantity and purity of
acetic acid as a top product.
The material chosen for the construction of the column is 304 stainless steel as both propionic and acetic
acid have corrosive properties and stainless steel provides a sufficient corrosion resistance to justify its choice
for the construction of the column. Grade 304 was chosen over other grades of stainless steel as the
mechanical and structural advantages provided by the other, more expensive, grades is not large enough to
justify the extra cost associated with them.
To calculate the internal diameter of the column a tray spacing has to be assumed, it is suggested that a tray
spacing of 0.5 m should be initially used to calculate the column’s diameter and if the diameter is greater than
1 m a tray spacing of between 0.3 and 0.6 m is normally appropriate (Sinnott et al., 2009, p708-709). The
calculated column diameter was greater than 1 m, so the initial assumed tray spacing on 0.5 m was carried
forward throughout the calculations. The tray efficiency applied to the tray in this case is 70% as this is found
to be an optimal number for the preliminary design of a distillation column (Sinnott et al., 2009, p700). It is
also suggested that 10% more tray be added in addition to tray efficiency with future expansion in mind
(Branan, 2005, p444).
In addition to the spacing of the trays and tray efficiency, included in the cost analysis of the column is the
costing of the trays used. This required a type of plate contactor to be chosen from a selection of sieve plate,
bubble-cap plate and valve plate, each with their own advantages and disadvantages. The plate contactor
chosen in this case is the valve plate, as there are weeping issues with sieve plates at low liquid flow rates
and bubble-cap plates are approximately twice as expensive as valve plates. The valve plate is an ideal
compromise between performance and cost when compared to the other two options.
Before the tray efficiency is applied to the theoretical number of stages, there are a total of 15 rectifying
stages and 8 stripping stages giving a total of 23 theoretical stages. The location of the column inlet stream

43
lies between the rectifying and stripping sections therefore, in this case the inlet stream is between the 15th
and 16th
stage, although this does not hold when the tray efficiency is applied as the number of stages
changes. Although the number of stages changes, the location of the feed should maintain the same ratio of
rectifying and stripping stages above and below it respectively, therefore, the initial ratio of rectifying stages
to total number of stages will be applied to the actual number of stages to find the actual inlet location of the
column.

44
4.4.2 Heavy – Ends Distillation Column P&ID
DC-501
P-501
P-502
V-502
V-503
V-504
V-505
FI
FIFIC
FIC
V-501
LIC
LI
V506
TI
FIC
V-507
TI FIC
Acetic acid
Propionic acid
LAH
LAL
Process
Steam
Process
Water
DC-401
V-508
FIC FI
Figure [4.11]: P&ID of Heavy-Ends Distillation Column

45
4.4.3 Acetic acid properties within the column
The vapour density of acetic acid under the conditions of the column was calculated using the ideal gas
equation (Perry, 1997, p2-355):
ρ 𝑎𝑐𝑒𝑡𝑖𝑐 = 𝑃
𝑅 𝑠𝑝𝑒𝑐𝑖𝑓𝑖𝑐 𝑇⁄
ρ 𝑎𝑐𝑒𝑡𝑖𝑐 = 1.84 kg/𝑚3
The vapour pressure due to acetic acid under the conditions of the column was calculated using the Antoine
equation and the parameters specific to acetic acid (Sinnott et al., 2009, p451):
𝑃𝑎𝑐𝑒𝑡𝑖𝑐 = 𝑒 𝐴−( 𝐵
𝐶+𝑇⁄ )
𝑃𝑎𝑐𝑒𝑡𝑖𝑐 = 136.45 𝑘𝑃𝑎
Where: A = 7.38782 (Dean, 2005, p539)
B = 1533.313 (Dean, 2005, p539)
C = 222.309 (Dean, 2005, p539)
The mass flowrates and mass fractions for the inlet, bottom product and top product have been calculated
using values taken from the simulation produced on Aspen Plus as displayed by the PFD:
Molar flowrate
[kmol/hr]
Mass flowrate
[kg/hr]
Mass fraction
Inlet 893.86 53631.48 0.992
Bottom product 3.07 184.17 0.296
Top product 890.79 53447.34 1.0
Table [4.8]: Acetic acid mass flowrates and mass flowrate for inlet, bottom product and top product.
4.4.4 Propionic acid properties within the column
The liquid density of propionic acid was calculated by developing a relationship between known values of
propionic acid density and temperature (CAMEO Chemicals, 1999) and assuming the relationship remained
constant up to the operating temperature of the column.
ρ 𝑝𝑟𝑜𝑝𝑖𝑜𝑛𝑖𝑐 = 882.39 𝑘𝑔/𝑚3
The vapour pressure due to propionic acid under the conditions of the column was taken from literature and
was given as (Clifford et al., 2004):
𝑃𝑝𝑟𝑜𝑝𝑖𝑜𝑛𝑖𝑐 = 59.86 𝑘𝑃𝑎

46
The mass flowrates and mass fractions for the inlet, bottom product and top product have been calculated
using values taken from the simulation produced on Aspen Plus as displayed by the PFD:
Table [4.9]: Propionic acid molar flowrates and mass flowrate for inlet, bottom product and top product.
4.4.5 Relative volatility
As this is a binary mixture, the following relationship between vapour pressures can be used to calculate
relative volatility (Branan, 2005, p450):
𝛼 =
𝑃𝑎𝑐𝑒𝑡𝑖𝑐
𝑃𝑝𝑟𝑜𝑝𝑖𝑜𝑛𝑖𝑐
𝛼 = 2.2795
4.4.6 Heavy Ends Distillation Column Specification
The final heavy ends distillation column specification is based on the calculations in Appendix E. The heavy
ends distillation column is represented diagrammatically in Figure 4.9.
The mechanical properties of the column were calculated based upon the Smoker equations which are
applicable to systems in which the relative volatility is close to 1 as the McCabe-Thiele method would be
impractical (Sinnott et al., 2009, p 661). Using this method and applying a 10% over design as well as a tray
efficiency of 70% give an actual number of stages required to be 37.
𝑁𝑢𝑚𝑏𝑒𝑟 𝑜𝑓 𝑎𝑐𝑡𝑢𝑎𝑙 𝑠𝑡𝑎𝑔𝑒𝑠 = 37
The height of the column is calculated by multiplying the tray spacing by the number of stages as well as the
addition of height allowances for a condenser at the top of the column and a reboiler at the bottom of the
column. The height allowance for the condenser and reboiler are suggested as 4 ft (≈1.25 m) and 6 ft (≈1.85
m) respectively (Branan, 2005, p 444). The column height, not including either end, is calculated to be 21.6
m.
𝐶𝑜𝑙𝑢𝑚𝑛 ℎ𝑒𝑖𝑔ℎ𝑡 = 21.6 𝑚
The column diameter is calculated as a function of the maximum vapour velocity through the column, which
is a function of tray spacing. The maximum vapour velocity through the column is calculated as 2.86 m/s and
the column diameter is calculated as 1.83 m.
Molar flowrate
[kmol/hr]
Mass flowrate
[kg/hr]
Mass fraction
Inlet 5.93 438.76 0.008
Bottom product 5.93 438.76 0.704
Top product 0 0 0

47
𝑀𝑎𝑥𝑖𝑚𝑢𝑚 𝑣𝑎𝑝𝑜𝑢𝑟 𝑣𝑒𝑙𝑜𝑐𝑖𝑡𝑦 = 2.86 𝑚/𝑠
𝐶𝑜𝑙𝑢𝑚𝑛 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 1.83 𝑚
The column shell thickness chosen is a calculated minimum shell thickness required to resist the internal
pressure with the addition of a corrosion allowance of 2 mm. This equates to a column shell thickness of 12
mm.
𝐶𝑜𝑙𝑢𝑚𝑛 𝑠ℎ𝑒𝑙𝑙 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = 12 𝑚𝑚
The thickness of the ends also has to be calculated as this will differ from the column shell thickness as the
stresses the ends of the column are put under vary from that of the column shell. The thickness of the end is
equal to 18 mm including a 2 mm corrosion allowance.
𝐶𝑜𝑙𝑢𝑚𝑛 𝑒𝑛𝑑 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠 = 18 𝑚𝑚
The size of the inlet, bottom product and top product pipes are all calculated using the density of the stream,
the velocity of the stream and the streams flowrate. The inlet pipe diameter was calculated to be 0.084 m
which corresponds to a nominal pipe size of 4 inches including an allowance for future expansion. The bottom
product pipe diameter was calculated to be 0.0092 m which corresponds to a nominal pipe size of 3/4 inches
including an allowance for future expansion. The top product pipe diameter was calculated to be 0.78 m
which corresponds to a nominal pipe size of 36 inches including an allowance for future expansion.
𝑖𝑛𝑙𝑒𝑡 𝑝𝑖𝑝𝑒 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 4 𝑖𝑛𝑐ℎ𝑒𝑠
𝑏𝑜𝑡𝑡𝑜𝑚 𝑝𝑟𝑜𝑑𝑢𝑐𝑡 𝑝𝑖𝑝𝑒 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 3/4 𝑖𝑛𝑐ℎ𝑒𝑠
𝑡𝑜𝑝 𝑝𝑟𝑜𝑑𝑢𝑐𝑡 𝑝𝑖𝑝𝑒 𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 = 36 𝑖𝑛𝑐ℎ𝑒𝑠
The cost of the column shell is a function of the columns mass which is calculated to be 120,000 kg. The total
cost of the column including the cost of the shell and the trays is £145,000 (2013, UK basis).
𝑠ℎ𝑒𝑙𝑙 𝑚𝑎𝑠𝑠 = 11914.6 𝑘𝑔
𝑐𝑜𝑠𝑡 𝑑𝑢𝑒 𝑡𝑜 𝑐𝑜𝑙𝑢𝑚𝑛 𝑎𝑛𝑑 𝑡𝑟𝑎𝑦𝑠 = £145,002
4.4.7 Summary of design data
To follow is a summary of the full design data of the column, calculations as to how the data has been
obtained is described fully in Appendix E.

48
Tray selection Valve plate trays
Tray spacing 0.5 m
No. of stages 37
Feed location Stage 24
Column height 21.6 m
Column diameter 1.83 m
Shell thickness 12 mm
End thickness 18 mm
Inlet pipe diameter 4 inches
Bottom product pipe diameter 2 inches
Top product pipe diameter 36 inches
Column mass 11914.6 kg
Installed cost of column shell and trays £145,000
4.4.8 Conclusion
The mechanical design of this unit has been completed with consideration of the 20% overdesign required
by the design brief as well as the product specification. To follow will be an engineering drawing that
describes all the data put forward, to give a clear visual representation of the design of the heavy ends
distillation column.

49
4.4.9 Engineering Drawing of Heavy – Ends Distillation Column
Figure [4.12]: Mechanical Design of Heavy – Ends Distillation Column.

50
4.5 Absorption column
4.5.1 Introduction
The production of waste gas components is inevitable in the case of carbonylation. The off was waste
materials (CO2, H2, CO) have to be separated from any toxic, and carcinogenic (e.g. methyl iodide)
components still present in the waste stream and then burned in the flare. It is necessary that the methyl
iodide is recycled back into the system, as it is highly toxic to the environment, and because it is required
as a reaction promoter for the carbonylation step.
The off gases are produced in the reactor phase as unwanted products are burned in the flare. While the
methyl iodide is captured in a counter-current packed absorption column using methanol and acetic acid
(i.e acetic acid is used for start-up, while methanol is sued throughout the life of the plant).
Absorption Column – Design Data
Height of transfer units 1 m
Height of packing section 8 m
Total height of column 10.5 m
Column Diameter 0.6 m
Pressure drop 0.005 bar g/m
Packing – Design Data
Type Intalox ceramic saddles
Size 25 mm (1 inch)
Packing material Ceramic
Packing arrangement Dumped

51
4.5.2 Absorption Column P&ID
Figure [4.13]: P&ID of Absorption Column
S-601

52
4.5.3 Design Method
Absorption is mass transfer procedure in which one or more soluble components from the gas mixture are
dissolved using a low volatility liquid. As a result the polluting material (i.e. methyl iodide) diffuses from a
gaseous sate into a liquid state, and is then recovered at the bottom of the column. The absorption rate is
driven by the driving force of the absorption, and is relatively independent of equipment used (McCabe and
Smith, 1976).
The absorption unit operates using a counter current design, where the methyl iodide present in the
gaseous mixture is dissolved in a liquid with a lower volatility (Sinnott et al., 2009). Counter current designs
have the highest theoretical removal efficiencies, and is suited for high loadings of pollutant materials while
it requires a lower solvent to gas ratios than alternative designs (e.g. crosscurrent, concurrent).
The most common choice for pollution control gas absorbers are packed towers. A packet tower is often
preferred to plate/tray towers because it can manage higher flowrates of gas, with lower pressure drops
while maintaining low liquid hold-up. It is also recommended to use a packed tower when the contacting
components have corrosive/acid proprieties, as cheaper corrosive materials are for the shell, and packing
are available. Also packed towers are preferred when there are no high temperature deviations, and the
system removes the gaseous mixture using a pressure drop. And when the diameter of the column (i.e.
based on the flowrates of material) is between 0.5-0.7 m (Sinnott et al., 2009).
Disadvantages of packed towers
• High clogging and fouling potential • Replacing damaged packing
• Higher waste water/solvent disposal • Removal of very small particles
4.5.4 Absorption Column Specification
The waste gas stream enters through the bottom of the column and travels vertically, counter current (i.e.
through the packing) to the falling solvent liquid. As a result, gaseous methyl iodide diffuses into a liquid
phase. The system works based on physical absorption, and achieves high efficiencies at low temperature
and pressure (Sinnott et al., 2009). Physical absorption is used because it relies entirely on the proprieties
of the solvent and the gas stream, and their specific characteristic (e.g. volatility, density, viscosity). In
order to achieve efficient absorption is it important that the design allows large contact area for the gas
stream and solvent to react, the capacity required for controlling high rates of waste gas, higher gas to
liquid ratios, low pressure drop and adequate distribution of solvent to gas to allow adequate pollutant
diffusion (i.e. methyl iodide)
4.5.4.1 Design considerations to account for drawback of unit
• A liquid distributor is used in order to maximize area covered by solvent in packing.
• A higher density component (e.g. acetic acid, methanol) is used as a scrubbing liquid.

53
• The solvent used is present in the system, and is recycled into the reactor with the methyl iodide
instead
• Packing material is corrosive-resistant, meaning that damage only occurs due to physical contact;
• Packing size is chosen based on the column size, so that it maximizes particle interaction, including
very small particles.
• Dumped packing allows easier replacement in the case of damage as opposed to structured
packing.
• Packing components that are not damaged can be reused, therefore reducing costs.
• A low pressure drop will results in low energy requirements.
4.5.4.2 Choice of packing
For this unit, the packing is the most important component. This is because the absorption efficiency is
correlated to the flow capacity, and the height of the transfer units (i.e. HTU gas, HTU liquid); these factors
significantly affect the tower height of the unit, and has economic implications (e.g. installation,
maintenance, cleaning). When choosing the adequate packing, it is important that the following factors are
taken into consideration:
• The packing material has to be inert to the liquids flowing through the packing.
• The packing material needs to be corrosive resistant.
• Brittleness of component needs to withstand process conditions, without presenting excessive
weight.
• The packing must offer enough contact area, while not restricting the gas and solvent flow.
• The packing must restrict the formation of excessive liquid hold-ups.
• The packing material needs to be acquired at a reasonable cost.
The above considerations were taken into account and 25 mm Intalox ceramic saddles were chosen.
These choice poses process advantages and low cost of packing. Intalox ceramic saddles offer the best
contact area, are inexpensive, and in ideal conditions could last throughout the life of the column. It is
expected that the packing will have to be changed throughout the life of the plant. However due to the small
costs, it still supports the choice of dumped packing rather than structured packing, or plates (Sinnot et al.,
2009).
4.5.4.3 Choice of absorption tower equipment
The packed absorption column is comprised out of:
• Column Shell
• Mist eliminator
• Liquid distributor
• Packing restrainer

54
• Packing support
• Packing materials
The mist eliminator is in the form of a layer of mesh; its main function is to collect any droplets that gather
at the top of the column as mist. It needs to be installed at the top of the column, so that any of the liquid
droplets collected are returned back to the column. The droplets are moved to the top of the column via a
high velocity gas stream (McCabe and Smith, 1976).
The liquid distributor chosen is a pressure drop spray nozzle distributer. The distributor is designed to wet
the packing, and facilitate consistent contact between the gas mixture and the solvent without constricting
the gas flow. Its main function is to spread the solvent evenly across the area of the packed bed. Some of
the disadvantages of this equipment includes plugging, formation of mist, feed rate dependent liquid
distribution. Therefore adequate maintenance is required in order to maximize the efficiency of the column
(Sinnott et al., 2009).
A packing support is necessary for an even distribution of the waste gas, and requires an open space
between the bottom of the absorption tower and the packing. The support plates are required to support the
total weight of the packing (i.e. while still allowing the material streams to travel freely), and are therefore a
necessity for this system (Sinnott et al., 2009).
A packing restrainer is required in order to prevent the high gas velocities from raising the packing into the
liquid distributors. The packing restrainer used is an unattached weighted plate placed at the top of the
packing, and which settles with the packed bed. The restrainer is required since the packing material is
ceramic, and keeping the integrity of the packed bed; therefore preventing any extra costs.
4.5.4.4 Materials of construction
The material of construction for the absorption column is Hastelloy C. Although, Hastelloy C is more
expensive when compared to a stainless steel shell lined with corrosion resistant column internals (e.g.
fibre-reinforced polymers), for this process route, Hastelloy C offers a better range for temperature and
pressure resistance. Additionally, the corrosion allowance for Hastelloy C accounts for extended use, with a
small probability of loss of containment across the column.
4.5.4.5 External equipment
The external equipment for the absorption column is comprised out of:
• Off gas movers
• Solvent pumps
• Control equipment

55
The waste gas movers require to be cooled down to room temperature, as increased gas temperature
leads to lower absorption in the column. Methyl iodide is gaseous above 315.9 K, and is required that the
methyl iodide is cooled down to 298 K. The resulting stream after cooling down will be a mixture of gas and
liquid, helping to absorb the methyl iodide as it turns into liquid phase, as well as keeping the vessel
temperature at room temperature in order to maximize the efficiency of the absorption column.
The solvent is moved into the column using a centrifugal pump. It is recommended that the construction
material used is corrosive resistant (e.g. stainless steel) and suitable for acetic acid/methanol. The
scrubbing material will be pumped from the final pure acetic acid stream. It is recommended that a storage
tank for acetic acid is placed in the proximity of the scrubber, in order to have access to excess scrubbing
solvent if the situation requires (Sinnott et al., 2009).
4.5.4.6 Safety control
The absorption unit requires the following control in order to operate safely on the premises of the plant:
• Gas detectors located at the outlet vent for: Acetic acid, Methanol, Carbon Monoxide, Carbon
Dioxide, Hydrogen, Hydrogen Iodide, Methyl Acetate, and Methyl Iodide.
• Temperature indication control for the gas stream.
• Level indication control for liquid stream.
• Flow indication control at the inlet and outlet streams.
4.5.4.7 Safety considerations
Since the components entering and exiting the absorption system are hazardous, adequate maintenance
for the equipment should be done regularly (HSE Maintenance procedures, 2002). A safe control
methodology has to be put in place. In order to keep the safe working environment the scrubber requires:
• Pressure relief system to prevent pressure accumulation in the vessel that could lead loss of
containment.
• Adequate insulation around the column with mineral wool to prevent any unwanted temperature
deviations.
• Gas detectors located around the scrubbing unit.
4.5.5 Conclusion
It is necessary that an efficient scrubbing system is put in place in order to prevent the release of toxic
iodide into the atmosphere. It is recommended that the above dumped packing absorption tower to be
connected to a stripping column, in order to achieve an efficient pollutant removal. The methyl iodide,
alongside the hydrogen iodide are toxic materials and have to be recycled back into the stream, or
disposed of adequately. These two components are crucial for the process to operate, and an efficient
pollutant removal system is necessary. When considering the design of the absorption column, a packed
absorber operating at low temperature and low pressure will offer the required removal efficiency. The

56
present design offers an economic alternative to other scrubbing units. The absorption unit is the safeguard
that prevents any hazardous components from polluting the environment heavily, with the adequate
safeguards and system management in place could last throughout the life of the plant.

57
4.5.6 Engineering drawing of Absorption Column
Figure [4.14]: Mechanical Design of Absorption Column

58
4.6 Storage tank – Acetic acid
4.6.1 Introduction
The tank is specified to be a fixed cone-roof cylindrical-type design and have a capacity of 8800 m3
. This
figure is the sum of one week production of acetic acid (7400 m3
) plus an extra 1400 m3
provided by the
volume of the conical head. The maximum working level for the tank is 6000 m3
, approximately 80% the
volume of the cylindrical tank and available for additional storage due to safety purposes. The storage
buffer provided will allow the plant for continuous operation for up to one week in the event of unforeseen
shutdown of the plant and/or disruption in distribution. The tank internal diameter is 24.9 m and the tank
height is 18.8 m. The internal and external pressure loads require a wall, base and roof plate thickness of
8.8 mm to meet the British design code for pressure vessels. This thickness gives a 100% safety factor
over the maximum anticipated stresses.
The product inlet line is standard nominal pipe size 4, schedule number 40s. This inlet is sized for the
maximum production flowrates. The product outlet line is of standard nominal pipe size 12, schedule
number 120. This line is sized such that a standard-size chemical ship tanker may be filled in 8 hours.
The tank must be constructed of stainless steel type 316L (‘acetic acid grade’), the specification of this
material is given in Appendix F. The design data required for this unit are specified below.
Acetic Acid Tank – Design Data
Design tank capacity 8800 m3
Design temperature 40 o
C
Design pressure 18.7 kPa
Working pressure 118 kPa absolute
Acetic acid density 1049 kg/m3
Material of construction SS316L
Design tensile strength 485 MPa
Joint efficiency 85%
The design tank capacity is estimated to be 8800 m3
. The maximum tank operating level will be
approximately 6000 m3
to give the extra tank capacity as a reserve volume. This also ensures that a
minimum of capital cost (in the form of product acid in the tank) is unused.
The design temperature represents the upper limit that acid may be fed to the tank from the process. The
working pressure represents the sum of atmospheric pressure and acid vapour pressure at the design
temperature. Details of the calculations associated with the tank design are presented in Appendix F. The
cost of this vessel is estimated (from a correlation) to be £333,000.

59
4.6.2 Acetic Acid Storage Tank P&ID
Figure [4.15]: P&ID for Acetic acid storage tank.
T-101

60
4.6.3 Design Method
The tank dimensions are determined according to standard tank geometries as enforced by API 650. Tank
shell thickness is sized according to the limitations imposed by the British design standard for pressure
vessels. The tank contents are flammable, toxic and corrosive, appropriate safety features are
recommended. Details of the calculations are given in Appendix G.
4.6.4 Storage Tank Specification
The final tank specification is based on the calculations in Appendix G is shown in Table 4.10. The tank is
represented diagrammatically in Figure 4.16.
The storage tank should be constructed of 316L stainless steel. To reduce corrosion of the tank bottom
exterior, application of coating is recommended between the tank and foundation. As seen by Figure 5.2,
the storage tank area is enclosed by containment facilities capable of containing the contents of the tank
and maximum expected rainfall in case of a storm event; moreover, additional safety of a small, deep diked
area is proven through lower evaporation rate and small area of fire. The enclosed area is drained through
a trap to a safe location that is protective of human health and environment and in compliance with
applicable laws and regulations. A vertical tank is implemented to provide for a more economical use of
land.
For outdoor storage of glacial acetic acid, a heating system and tank insulation is provided. The
recommended heating system consists of low-temperature electric heating pads installed between the tank
exterior and the insulation in order to maintain the temperature at a desirable level.
Acetic acid is a flammable solvent, thus to inhibit the accumulation of static charges, the storage tank,
pumps, transfer lines, and offloading vehicle are adequately grounded and fill line enters the tank through
the roof and extended downward to within 2 or 3 inches of the bottom.
The storage facility is constructed so that water cannot be introduced or generation of heat occurs. In a
confined space, considerable pressure caused by this reaction can result in an explosion that may rupture
the storage tank.
Safety features for the tank include a pressure relief-valve system on the tank roof, to be opened when
draining or adding to the tank contents. An emergency relief vent is fitted to the storage tank to allow
emergency flow due to excessive venting requirement from fire burning around the tank, thus eliminating
opportunity for a costly tank rupture, providing emergency venting from abnormal internal pressure beyond
the capability of the pressure relief vent. The operational tank venting system handles normal tank venting
due to product import/export and ambient temperature variations. In the event of fire, as vapour pressure
increases to a point where normal venting equipment capacity is exceeded, the hinged cover will lift
relieving the pressure and protecting the tank from rupture. The pressure build up will be quite slow,
therefore the cover should not open violently and cause any damage to the tank. Emitted vapours may be
ignited by the fire, but should ‘flame off’ externally until brought under control by firefighting operations.. A
manhole provides access to the tank for internal maintenance.

61
Appropriate venting systems are issued. Vents should be angled at 45o
from vertical and cut off vertically to
prevent rain from entering. The vents are 1 inch larger in diameter than the tank fill line. A coarse-mesh
stainless steel wire screen is placed over vent openings to prevent entry of foreign objects. A blanket of
inert gas, nitrogen, is provided and equipped with a pressure/vacuum conservation vent, piped away to a
safe location that is protective of human health and the environment.
In order to counteract the strong odour generated by acetic acid, odour masking methods are utilised within
the compound. Odour control chemicals used for masking purposes are aromatic chemicals derived from
aromatic chemical manufacture. Organic odour control chemicals are numerous, some of which are vanillin,
methyl ionones, benzyl acetate, phenylethyl, eugenols and heliotropin, and thus any of the organic
compounds mentioned are viable to be implemented in this scheme. Furthermore, this method of treatment
requires no capital investment required for equipment and readily available for application.
Table [4.10]: Storage tank specification.
A standard metal staircase and railing skirts the outer edge of the tank providing access to the tank roof. A
manhole in the tank roof provides access for internal repairs. Discrete inlet and outlet lines are required to
feed into the base of the tank. A pressure relief valve is attached to the roof. This valve is opened
automatically when pumping product to, or withdrawing product from the tank. The valve is shut when
pumping stops so that vapour losses from the tank are contained. A bursting disc on the roof also provides
emergency pressure relief for an unforeseen pressure build-up within the tank.
4.6.5 Conclusion
The acetic acid storage tank required for the plant has a capacity of approximately 8800 m3
. This
represents approximately 21,000 tonnes of product acid. The tank will normally contain only about 14,000
Vertical cylindrical-type tank
Fixed conical roof
Total tank capacity 8800 m3
Normal operational capacity (maximum) 7400 m3
Tank inside diameter 24.9 m
Tank height 18.8 m
Tank wall thickness 8.8 mm
Material of construction SS316L
Inlet and outlet at tank base
Inlet line: Nominal pipe size 4, schedule number 40s
Outlet line: Nominal pipe size 10, schedule number 120
Manhole, bursting disc, pressure relief valve on roof
Mineral wool insulation 0.02 m

62
tonnes to satisfy outside product sales. The remaining capacity is in reserve in case of plant shut-down and
issues regarding distribution. This provides a one week production buffer for the plant. The final
specification requires a tank of 24.9 m diameter, height 18.8 m, and a plate thickness of 8.8 mm.
A size delivery of 5500 m3
by means of a standard 5000 tonne chemical tanker would be desirable, as
nearly 75% of the weekly production is able to be shipped. The figure obtained for the minimum stock level
will be able to accommodate a scenario where if the plant shuts down for a few days, the remaining storage
will keep on supplying customers. Additionally, if there is a problem in distribution, the available space in
the tank will allow for continuous running of the plant until the problem is solved. Henceforth, a delivery
period of one week will provide smooth production and shipment whilst accommodating customers. The
cost of the tank has been calculated (Appendix G) from correlations to be approximately £333,000.

63
4.6.6 Engineering drawing of Acetic Acid Storage Tank
Figure [4.16]: Mechanical design of Acetic acid storage tank.

64
5. Process Control and Instrumentation
5.1 Introduction to Process Control and Instrumentation
The premise of process control and instrumentation consists of applying the philosophies of control to all
aspects of a process, whether that be the design of the process, the standard operation of the process or the
operation of the process under conditions that vary from the norm, for example during a start up or shut down,
as the only way to have a safe system, is to have reliable control systems in place at all times. As well as
safety, putting in place reliable and effective control systems also aids in improving plant efficiency and the
economic stability of the process while ensuring the plants compliance with the relevant environmental and
safety regulations.
As the philosophies of process control and instrumentation are incorporated into every aspect of a process,
the instrumentation that’s implemented into a process should be integrated within the earliest design instead
of being an afterthought that’s ‘bolted on’ once the design has been completed. Implementing process control
systems in this way will lead to the safety features of the process being much more effective with regards to
plant safety and the safety of those in and around the plant, in addition to being much simpler because they
have not been worked into the system, around an already complete process without the necessary control
features put in place.
5.2 Objectives of Process Control
The primary objective of the control systems and instrumentation put in place is to ensure that the process is
carried out in a safe and reliable manner while producing a product that meets the required specification
desired by the consumer. This can be achieved by (Sinnott et al., 2009, p275):
1. Maintaining process variables within known safe operating limits around a specified set point.
2. Alerting operators to deviations from the set point and safe operating limits and provide a solution to
the deviation either via the manipulation of process equipment or shutdown systems.
3. Preventing operators from altering process variables such that operation outside the safe operating
limits is caused.
The four main process variables that have been controlled throughout the design stages of this project are:
1. Temperature.
2. Pressure.
3. Level.
4. Flow.
The variables have been controlled through the manipulation of various pieces of process equipment, such
as control valves, via information sent to them through a series of indicators, alarms and controllers. The
process equipment works to control any deviations from the desired set point, within an upper and lower limit.
The control measures also have to accommodate for planned changes in the set point, for example, during

65
start-up and shut down in which the desired set point for each of the process variables is going to vary from
that of a steady state operation. This can be seen in the reactor, for example, during its start-up process, the
set point would be at its desired operating conditions to ensure that the reaction took and a product was
being produced. As the reaction continues and begins to reach a steady state the heat released by the
reaction will cause an increase in the temperature of the reactor and therefore, a deviation from the desired
operating conditions. In this case the set point can be decreased below the desired operating conditions while
the heat released by the reaction can be used to maintain the required temperature. Variations in the set can
be fine-tuned using the instrumentation in place to control situations like this.
In addition to process safety, operability has to be taken into consideration when implementing control
features throughout the design stages of a project, as having too many indicators and alarms, etc., can reduce
the operability of the process as having “too much data being thrown at operators reduces their ability to
understand what is happening and respond correctly” (Hurley, 2016). This has been taken into consideration
when implementing control features into the design of our process as it is possible to manage a combination
of process variables by manipulating just one of them. Although this could lead to undesired variations in
some of the process variables when manipulating others, therefore it is important to understand the
relationships each of the process variables share with each other and what impact their variation would have
on the process streams as well as the process equipment. An example of this can be seen in the transition
from stream 4 to stream 5 of the PFD through a pressure reduction valve. The purpose of the valve is to
reduce the streams pressure before it enters the flash tank, but as a consequence of reducing the pressure,
the temperature of the stream is also reduced. The relationship between the two process variables is clear
in this case, which would make it easier to identify the hazards a variation of pressure would cause
downstream of the valve whether that be to the flash tank or other pieces of process equipment.
5.3 Implementation of Control Systems in our Design
The types of control that have been taken advantage of throughout the design process of this project has
been feedback, feedforward and cascade, with feedback and feedforward being the simpler forms of control
and cascade being more complex. A feedback control system measures a process variable downstream of
a piece of process equipment and then send information back upstream for the process equipment to manage
the process variable directly. Similarly to feedback control, feedforward control only measures and alters one
process variable, although the process equipment is downstream of the measuring equipment instead of
being the other way round in feedback. Whereas cascade control measures a process variable then alters
another to result in a change to the original measured process variable.
Feedback control can be seen throughout the process P&ID, an example of the utilisation of feedback control
can be seen in figure 6.1. The figure shows a set of pumps in parallel with feedback controlled valves based
on the flow through the pipe. This has been implemented into the system with pump failure in mind, for
example, if P-401 was the operational pump and P-402 was being used as a backup, both V-402 and V-404
would be closed to prevent any flow through P-402 under normal operating conditions. In the event of P-401
failing, V-403 and V-405 could be closed from the control room to prevent flow through P-401 while the pump

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