School of Chemical and Process Engineering
CAPE3000 DESIGN PROJECT (2023-24)
PART-1 (Group) REPORT, Semester 1
Carbon Capture Plant
GROUP NUMBER: CEE2
GROUP MEMBERS: Faisal Kamal, Abdelrahman Abouelela, Ahmed Aldhalei,
Bryn Barker, and Sari Miari
Company Name: FAABS
Writing to the board of Australian Energy (AE) Engineering Ltd
SUPERVISOR: Timothy Cockerill
SUBMISSION DATE: 14 December 2023
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
DECLARATION OF INDIVIDUAL CONTRIBUTIONS
Chapter
Lead author(s)
[% contribution]
Signature(s)
Sari Miari (85%)
Sari
Sari Miari (70%)
Sari
Chapter 2: Process
Route Evaluation
& Selection [10%]
Chapter 3: Design
of Selected Process
[15%]
Chapter 4: Raw
Materials &
Products [5%]
Chapter 5: Safety,
Health &
Environmental
[15%]
Chapter 6: Mass &
Energy Balances
and PI [50%]
Other parts of the
report (executive
summary,
introduction,
conclusions,
nomenclature,
references) [5%]
Other members
contributed to this
chapter.
[% contribution]
All members are
required to contribute to
this chapter
Abdelrahman Abouelela
(5%)
Faisal Kamal (5%)
Bryn Barker (5%)
All members are
required to contribute to
this chapter
Signatures
Abdelrahman
Faisal
Bryn
Faisal
Bryn
Faisal Kamal (25%)
Bryn Barker (5%)
Abdelrahman
Abouelela (90%)
Abdelrahman
Sari
Faisal Kamal (70%)
Faisal
Bryn Barker (30%)
Faisal
Bryn
Plant Calculations –
Bryn Barker (80%)
DCC – Sari Miari
(100%)
Absorber – Bryn
Barker (95%)
Stripper – Faisal
Kamal (95%)
HEX – Bryn Barker
(98%)
Compressor Train –
Abdelrahman
Abouelela (100%)
Pinch Point – Bryn
Barker (100%)
Abdelrahman
Abouelela (75%)
Bryn
Abdelrahman
Faisal
Sari
All members are
required to contribute to
this chapter
Bryn
Abdelrahman
Faisal
Sari
Sari Miari (5%)
Ahmed Aldhalei (5%)
Plant Calculations –
Abdelrahman Abouelela
(20%)
Absorber – Faisal
Kamal(5%)
Stripper – Bryn
Barker(5%)
HEX – Ahmed Aldhalei
(2%)
Abdelrahman
ii
Sari Miari (25%)
Sari
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
EXECUTIVE SUMMARY
Following an extensive process route selection and evaluation, post-combustion carbon capture,
coupled with MEA as the amine solution for absorption, has emerged as the most optimal strategy
for this project. MEA's increased absorption capabilities, stemming from its reactivity, makes it
optimal for our objectives.
To maximise cost-effectiveness, we propose sourcing construction raw materials from Chinese
suppliers, while also securing coal bed methane from China. Conversely, due to market stability
and supplier flexibility, MEA from a local source is recommended.
Operational safety is paramount, given the potential hazards of certain chemicals used in
operation. Rigorous measures, including exposure controls and proper storage and disposal
protocols, are put in place to mitigate potential catastrophe, such as wastewater undergoing
treatment to avoid adverse environmental impacts. The operational units in the process have
consistently met benchmarks, with the absorber column achieving over 90% CO2 absorption from
the flue gas and the stripper generating a CO2 stream exceeding 95 wt% purity.
Following compression and being liquefied, the CO2 stream is sent to an underground pipeline to
be stored. Calculations project ≈13 kg/s of CO2 to be stored at full operational capacity. The
strategic decision to store rather than sell the captured CO2 to collect Australian Carbon Credit
Units (ACCUs) is heavily advised. The project's anticipated capture and storage of 370,000 tons of
CO2 annually, valued at over $4.5 million, not only ensures economic viability but also exhibits our
commitment to positively impact the environment by reducing CO2 emissions.
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
ACKNOWLEDGEMENTS
Thank you to Tariq Mahmud, Timothy Cockerill and Manoj Ravi for all help and advice provided to the
group regarding the production of this report.
This report contains Abdelrahman Abouelela, Bryn Barker, Faisal Kamal, Sari Miari and Ahmed Aldhalei’s
own unaided efforts.
Signatures:
Abdelrahman
Faisal
Bryn
Sari
iv
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Table of Contents
1. INTRODUCTION ........................................................................................................................................................................ 1
2. PROCESS ROUTE EVALUATION AND SELECTION ........................................................................................ 2
2.1 Processes Route Evaluation ...................................................................................................................................... 2
2.2 Process Selection ................................................................................................................................................................. 3
2.2.1 Technological Methods for Capturing CO2 ..................................................................................................... 3
2.2.2 Choice of Absorbent ............................................................................................................................................. 4
2.2.3 MEA Degradation Products............................................................................................................................... 6
2.2.4 Chemicals physical properties ......................................................................................................................... 7
2.2.5 Simplified Process Pathway including Main Process Units .............................................................. 7
2.3 Plant Location .................................................................................................................................................................... 7
2.4 Environmental Impact .................................................................................................................................................... 9
3. Design OF Selected Process ......................................................................................................................................... 10
3.1 Process Description & Flowsheet Design ........................................................................................................ 10
3.1.1 Pre-treatment Units ............................................................................................................................................ 10
3.1.2 Direct Contact Cooler (DCC-101) ............................................................................................................... 11
3.1.3 Absorber Column (A-101) ............................................................................................................................... 11
3.1.4 Heat Exchanger and Cooler (H-101) ......................................................................................................... 12
3.1.4 Stripper Column (S-101) .................................................................................................................................. 13
3.1.5 Compressor Train (COM-101 TO C-105) ............................................................................................... 13
3.1.6 Process Flow Diagram...................................................................................................................................... 15
3.2 Process Chemistry & Thermodynamics ............................................................................................................ 16
3.2.1 Absorption of CO2 by water ............................................................................................................................ 16
3.2.2 Absorption of CO2 by MEA Kinetics & Mechanisms .......................................................................... 17
3.2.3 Whitman Two-Film Mass Transfer Theory ............................................................................................. 18
3.2.4 Vapor-Liquid Equilibrium (VLE) Data ....................................................................................................... 19
3.2.5 Side Reactions Kinetics and Mechanisms.............................................................................................. 22
4. RAW MATERIALS AND PRODUCTS ....................................................................................................................... 24
4.1 Raw Materials ................................................................................................................................................................. 24
4.1.1 Raw Materials for Construction: ................................................................................................................... 24
4.1.2 Raw Materials for Operation: ......................................................................................................................... 24
4.1.3 Coal bed methane supplier: ........................................................................................................................... 25
4.2 Product and By-products .......................................................................................................................................... 26
5. SAFETY, HEALTH, AND ENVIRONMENT ............................................................................................................. 26
5.1 Process Safety ............................................................................................................................................................... 26
5.2 Health ................................................................................................................................................................................. 28
5.3 Environmental Impact ................................................................................................................................................. 31
6. MASS & ENERGY BALANCES and Process Integration ................................................................................ 32
6.1 Mass Balance PBD ...................................................................................................................................................... 32
6.1.1 Flue Gas................................................................................................................................................................... 32
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6.1.2 Direct Contact Cooler (DCC-101) .................................................................................................................... 35
6.1.5 S-101 .................................................................................................................................................................................... 40
6.1.5.1 Condenser mass balance: ............................................................................................................................... 41
6.1.5.1.1 Stream 6: ......................................................................................................................................................... 41
6.1.5.1.2 Stream 7: ......................................................................................................................................................... 42
6.1.5.1.3 Stream 8: ......................................................................................................................................................... 43
6.1.5.1.4 Flowrates around condenser:................................................................................................................ 44
6.1.5.2 Mass Balance around Reboiler and column: .......................................................................................... 45
6.1.5.2.1 Stream 11 ........................................................................................................................................................ 45
6.1.5.2.2 Stream 10 ........................................................................................................................................................ 46
6.1.5.2.3 Stream 9 .......................................................................................................................................................... 46
6.1.5.2.4 Flow rates around the reboiler and column. .................................................................................. 47
6.1.5.3 Mass balance summary table......................................................................................................................... 48
6.1.6 Compressor train (COM-1 to COM-4, C-102 to C-105, D-101&D-102) ............................................ 49
6.2 Energy Balance ................................................................................................................................................................ 50
6.2.1 Direct Contact Cooler (DCC-101) .................................................................................................................... 50
6.2.2 Packed Bed Absorption Column (A-101)..................................................................................................... 51
6.2.3 - E-101: ....................................................................................................................................................................... 53
6.2.4 Cooler (C-102)......................................................................................................................................................... 54
6.2.5 S-101 Energy balance ................................................................................................................................................. 55
6.2.5.1 Condenser Duty .................................................................................................................................................... 55
6.2.5.2 Energy balance around the column............................................................................................................. 56
6.2.5.3 Reboiler Duty .......................................................................................................................................................... 57
6.2.6 Compressor train (COM-1 to COM-4, C-102 to C-105, D-101&D-102) ............................................ 58
6.3 Process Integration ........................................................................................................................................................ 62
6.3.1 Hot and cold utility requirements .................................................................................................................. 62
6.3.2 Selection of streams for PI ................................................................................................................................. 62
6.3.3 Minimum approach temperature ................................................................................................................... 62
6.3.4 Application of PI and design of HEN .............................................................................................................. 63
....................................................................................................................................................................................................... 65
6.4 Final Process Flow Diagr ............................................................................................................................................. 65
6.5 Auxiliary unit power requirement calculations ................................................................................................ 66
7. CONCLUSIONS .................................................................................................................................................................... 67
8. References ................................................................................................................................................................................. 68
APPENDIX B - Raw Materials and Products appendices: .............................................................................................. 79
APPENDIX C - Mass and Energy Balance appendices ......................................................................................... 81
APPENDIX D - MINUTES OF MEETINGS WITH SUPERVISOR: ........................................................................ 86
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
NOMENCLATURE
Symbol
CP
Quantity
Specific heat capacity
SI Unit
j/mol
Tc
Pc
Critical temperature
Critical pressure
K
bar
W
Work
Density
Enhancement factor
Enthalpy
Henrys Law Constant for species i
Molar Flow Rate
Molar Gas constant
Quantity of heat
J = kg m2 s-2
Kg m-3
ρ
E
̂
𝐻
ki
𝑛̇
R
Q
J = kg m2 s-1
Mol/L.bar
Mol s-1
J mol-1 K-1
J = kg m2 s-1
Greek letters
Symbol
Quantity
γ
Specific heat ratio
SI Unit
J/Kg K
Dimensionless Groups
Notation
xi
yi
zi
Ki
Name
Formula
Liquid molar fraction of component i
Vapour fraction of component i
Inlet molar fraction of component i
Equilibrium constant of component i
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ABBREVIATIONS
Coal Seam Gas – CSG
Coal Bed Methane – CBM
Carbon Dioxide – CO2
Monoethanolamine - MEA
Greenhouse gas - GHG
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
1. INTRODUCTION
This report aims to explain the process of collecting and storing CO2 from the flue-gas of a 100 MWe
gas-fired power station, based in Queensland, Australia. The power station utilises Coal Seam Gas
as its chosen fuel source, which is mainly composed of alkanes (99.5 vol%), with the majority of the
fuel being composed of methane (95 vol%). The building, design and optimisation of the process
has been discussed in detail, along with the appropriate uses for the captured CO2. A rigorous
process route evaluation and selection process was conducted to determine the most suitable
combustion process and absorbent used in operation for the most effective method of carbon
capture. In this process, details stretching from the absorbent’s physical properties to the chosen
plant location have been scrutinised in an effort to function as efficiently as possible while mitigating
environmental impacts that could arise from the carbon capture project.
Furthermore, a thorough description of the entire process has been provided, explaining the function
of every unit in detail, along with relevant data on any chemicals and theories that contribute to the
function of the units. The suppliers for the raw materials required for the whole carbon capture
process, from building the carbon capture unit to when it is fully operational, are examined and the
most suitable supplier is chosen, whether it is a local or international supplier. The potential uses of
the captured CO2 are discussed and the most appropriate use is chosen based on economic
impacts to FAABS as well as environmental impacts. The potential health and safety issues that
come along with building and operate the unit along with how to mitigate them have been discussed
in this report, to ensure all involved are safe throughout the entire process. Moreover, a rigorous
mass and energy balance has been conducted on every operational unit as well as the entire
process to understand the chemical and physical side of carbon capture. Process integration was
considered throughout to allow for maximum efficiency while also minimising expenses to make the
project as economically feasible as possible.
After designing the carbon capture plant and ensuring plant operation is feasible, the individual
components are to be designed. FAABS has compiled recommendations to the process to be used
based on the research and calculations conducted, which will be displayed in this report in an
attempt to establish the economic viability of the carbon capture project.
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
2. PROCESS ROUTE EVALUATION AND SELECTION
2.1 Processes Route Evaluation
As Coal Seam Gas (CSG) is fed into a gas turbine along with air, the CSG combusts to generate
electricity resulting in the release of a flue gas which contains CO2 which is a greenhouse gas that
has a negative impact on the environment due to its contribution to global warming. There are three
primary routes for capturing CO2 from flue gas emissions. These are oxy-fuel combustion, precombustion, and post-combustion.
Oxy-Fuel combustion carbon capture is a process wherein the bulk nitrogen in the air feed is
removed prior to combustion, yielding a stream that composes of ~95% O2; CSG combusts at a
higher temperature in this condition compared to air-fired combustion, increasing the efficiency of
the combustion process (Oxy-combustion, 2023). Moreover, the flue gas produced primarily
composes of CO2 and water vapour, which reduces the volume of the flue gas by approximately
75% compared to other methods, which leads to less heat lost in the exhaust, making the system
more energy efficient (Pronobis, 2020). Moreover, the concentration of CO2 in the flue gas is
significantly higher compared to using other methods, which simplifies the CO2 capturing process in
the carbon capture unit, as there are fewer other gases to separate it from, because the energy
required for CO2 separation is proportional to its concentration in the flue gas, and higher
concentrations of CO2 reduce the energy needed for its capture (Karimi et al, 2023). This process
has a potential for 100% CO2 capture (Basile et al, 2019). Moreover, the formation of NOx will be
avoided to a certain extent due to the lack of nitrogen in the gas turbine. However, an NOx treatment
unit might still be needed, as the air separation unit cannot achieve 100% oxygen purity and the
CSG may contain small impurities of nitrogen. Thus, little amounts of NOx will inevitably form.
Therefore, FAABS does not recommend that AE uses oxy-fuel combustion because purifying the air
stream to achieve an O2 rich stream is extremely expensive and energy intensive, and yet the system
could still require an NOx treatment unit. Furthermore, the increase of combustion and capturing
efficiencies do not overturn the increase of costs caused by purifying the air stream.
Figure 2.1: Simple process block diagram showing oxy-fuel Combustion CO2 Capture (Drawn by author).
Moving on, in pre-combustion capture, CO2 is captured prior to the fuel combustion. This multi-stage
process begins with gasification, where the CSG and steam at 750°C are fed along with an oxygen
rich stream into a gasifier (Paul, 2004), which transforms the CSG into syngas which consists of
carbon monoxide (CO) and hydrogen (H2) (Persons, 2019). Then the syngas along with steam enter
a Water-Gas Shift Reactor which converts the CO into CO2 (Persons, 2019), the CO2 is then
captured in a carbon capture unit. This separates the H2 from the CO2, so finally the purified H2 is
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
directed into the gas turbine where it is burned as a fuel to generate electricity (Persons, 2019). The
Pre-Combustion Carbon Capture method captures 90-95% of CO2 in flue gas (Basile et al, 2019).
This is because the stream that enters the carbon capture only contains CO2 and H2, which means
that the CO2 concentration in the stream is high which makes the capturing process more efficient.
However, FAABS does not recommend AE to use pre-combustion due to its high capital cost, and
due
to
the
complexity
of
the
process
and
the
need
for
specialised
equipment.
Figure 2.2: Simple process block diagram of Pre-Combustion CO2 Capture (Drawn by author).
Finally, in post-combustion capture, CO2 is captured from the flue gas simply after the fuel combusts
with air in the gas turbine. The post-combustion capture method captures around 90% of CO2 in flue
gas. Currently, post-combustion is the most established and commonly used approach due to having
the highest technology maturity and cost-effectiveness, and most importantly due to having the
lowest capital cost and indirect costs compared to the other methods (Theo et al, 2016).
Furthermore, post-combustion can be retrofitted onto existing power plants and industrial facilities
easily. This plays a major role in reducing the cost of adopting a carbon capture method into an
existing plant, it also allows the carbon capture unit to be replaced or upgraded with minimum to no
impact on the retrofit plant (Carbon capture, 2022). Therefore, FAABs advocates the adoption of this
method by AE on the grounds of its simplicity, cost-effectiveness, and low capital costs.
Figure 2.3. Simple Process Block Diagram of Post-Combustion CO2 Capture (Drawn by author).
2.2 Process Selection
2.2.1 Technological Methods for Capturing CO2
Having chosen to capture the CO2 from the flue gas after combustion with air has taken place, it is
important to consider the technology that shall be used to capture the CO2. The methods which have
been considered for the Plant in Queensland can be found in the table below:
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Table 2.1: Different possible technologies to use in capturing CO2.
CO2 Capture
Technology
Description
Advantages
Disadvantages
Absorption
A gas stream is contacted with a liquid absorbent
(solvent), absorbing CO2 either chemically (e.g.
Amines) or physically (e.g. Organic molecules). Then
heat and/or pressure are applied to release the CO2
and the absorbent is recycled in the system (Co,
2019).
Involves the intermolecular forces between the CO 2
and the surface of the adsorbent, resulting in CO 2
adhering to the surface. Heat, electricity, or pressure
is then applied to release CO2 (Co, 2019). The most
used adsorbent materials are activated carbon,
zeolites, and silica (Dziejarski et al, 2023).
A selective semi-membrane is used to separate the
flue
gas components
through
various
mechanisms such as diffusion, molecular sieve, and
ionic transport. The partial pressure difference across
the membrane provides the driving force for
separation (IEAGHG, 2019).
The flue gas is cooled by an external refrigerant loop.
Flue gas is initially cooled to remove water. Then
further cooling is achieved via direct contact with a
cryogenic fluid, forming a slurry with solid CO2. The
solid CO2 is then removed via filtration and warmed to
turn into liquid CO2 (Sayre et al, 2017).
Most
mature
technology
(Leung et al,
2014).
High energy intensity and solvent might
degrade over time. It also requires a
significant amount of water (Tuinier et
al, 2011).
Achieves 9095%
CO2
capture
(IEAGHG,
2019).
High energy intensity. Moreover, the
adsorbent can degrade or become
fouled, reducing efficiency. The
process is slow and can’t maintain
rapid CO2 capture (Hudson &
Thomson, 2019).
Removal of CO2 from flue gas is
difficult due to low CO2 partial pressure
and precise of water vapour, achieving
around
80%
carbon
capture
(Mohammed et al, 2020).
Adsorption
Membrane
Cryogenic
Low
energy
requirements
(Abanades et
al, 2015).
Can
achieve
CO2
capture
between 95
and
99%
(Baxter et al,
2018).
High energy intensity, with a very high
capital cost and many technical
difficulties. It is still under development
to improve the technology and make it
more economically viable (Aneesh &
Sam, 2023).
AE’s carbon capture plant in Queensland requires to remove 90% of the CO2, which excludes the
option of using the membrane technology as it only achieves around 80%. Moreover, the cryogenic
technology can achieve a greater carbon capture rate than what is required. However, the
technology is still under development as it is not economically viable yet. Therefore, the choice lies
on either absorption or adsorption, as both techniques can achieve the required carbon capture rate,
but both techniques are highly energy intensive. However, currently absorption is shown to be more
effective from an economic and operational point of view; with a larger processing capacity on the
industrial scale, with a better return and a better long-term performance (Castro et al, 2021).
2.2.2 Choice of Absorbent
As mentioned in table 2.1, absorption of CO2 from a flue gas can be done either chemically (e.g.
Amines) or physically (e.g. Organic molecules). Current data and research suggest that chemical
absorption is the most efficient option if the composition of CO2 in the flue gas is below 30.4% (Zhang
et al, 2020). This is because chemical absorption favours low CO2 partial pressures in the flue gas,
(Mumford et al, 2015). Based on calculations in section 6.1.1, AE’s proposed plant flue gas will
compose of 8.5 mol% CO2. Therefore, FAABS advocates the use of an amine solvent for chemical
absorption as it would be the most cost & energy-efficient option.
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The most widely used amines in the absorption of CO2 are monoethanolamine (MEA) and
diethanolamine (DEA) (Varbanov et al, 2017). MEA & DEA similarly react with CO2 in an aqueous
solution. Firstly, the carbon atom of CO2 binds with the nitrogen atom of MEA via a nucleophilic
attack, forming a Zwitterionic intermediate (MEA+COO-) (Holmes, 2018). Then H2O molecules cause
the deprotonation of MEA+COO-, forming a stable carbamate ion (MEACOO-) and a solvated proton
(H3O+) (Hwang, 2015). The H3O+ reacts with another MEA molecule to form MEAH+ and H2O, which
means that theoretically 2 moles of MEA absorb 1 mole of CO2 (Hwang, 2015). The mechanisms of
all these reactions are shown in figure 10 in section 3.2.2.
The loaded solvent, saturated with CO2 is then regenerated by applying heat to it as this reaction is
reversible, the carbamate turns back into the Zwitterion molecule which then dissociates into MEA
and CO2, releasing the CO2 from the solvent as a gas to be captured (Lv et al, 2015), while the
regenerated solvent can be reused in the absorption step.
MEA & DEA absorb CO2 optimally at 40-50 °C because the reaction is exothermic (Bravo et al,
2021). MEA is regenerated optimally at 110-130 °C (Bravo et al., 2021) and DEA is regenerated
optimally at 95-105 °C (Zhou et al, 2017). Furthermore, the heat of reaction of MEA is 85.6 kJ/molCO2
which is greater than needed for DEA 76.3 kJ/molCO2 (Galindo et al., 2012). Both of these factors
lead to a much higher energy requirement for stripping for MEA than DEA.
Moreover, DEA is less reactive than MEA, meaning that it has a lower chance of degrading/reacting
with impurities present in the flue gas (Xue et al, 2018), which means that DEA is less likely to
degrade overtime compared to MEA, and thus DEA is less likely to cause the corrosion of equipment.
However, a major disadvantage of DEA is that it exhibits slow kinetics; making MEA a better choice
as it can remove CO2 from the gas stream much more rapidly (Galindo et al, 2012). The faster
kinetics of MEA can be observed from comparing figures 4 and 5, as MEA takes noticeably much
less time to reach the equilibrium. This leads to a higher CO2 loading rate by MEA, meaning that
MEA can capture more CO2 per unit of solvent in a short period of time (Xue et al, 2018). This is
particularly beneficial in industrial processes, because the equipment used for the absorption
process can potentially be smaller. This leads to cost savings in terms of equipment size, installation,
and operational costs.
Figures
2.4 & 2.5: Equilibrium loading curve for MEA & DEA (Galindo et al, 2012).
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In Australia, MEA costs $1596/t, nearly the same as DEA $1546/t (Redox, 2023). Therefore, after
careful consideration of all factors, FAABS recommends the use of MEA as the absorbing solvent.
Furthermore, 30 wt% MEA in water is optimal for CO2 capture when the composition of CO2 (8.5
mol%) in the flue gas is low, this is due to the amine’s reactivity and absorption characteristics (Zhang
et al, 2020), which is discussed in section 3.2.4.3. Moreover, this balances the cost of MEA with the
increase in driving force for mass transfer for the absorption of CO2.
2.2.3 MEA Degradation Products
MEA can degrade overtime due to oxidative or thermal degradation, resulting in solvent loss and
production of harmful by-products which can cause equipment corrosion. Oxidative degradation
occurs when MEA reacts with O2, SO2 or NOx in the flue gas. The degradation of MEA by SO2
ultimately causes the formation of Bicine which can reduce the efficiency and performance of the
absorbing system, and cause equipment corrosion and environmental issues (Fytianos et al, 2015).
Furthermore, presence of NOx can cause the degradation of MEA by photo-oxidation to form
nitrosamines and nitramines, which can cause equipment corrosion and have also been identified
as carcinogenic and highly hazardous for human health and the environment (Vega et al, 2014).
Finally, oxidative degradation can also occur in the absorber if the flue gas O2 composition is around
3% or higher (Vega et al, 2014). The main products of MEA oxidative degradation are ammonia,
formic acid, acetic acid, HEF, HEI, and HEPO (Vega et al, 2014). These can reduce the efficiency
and performance of the absorbing system, and cause equipment corrosion.
Thermal degradation also inevitably occurs in the stripper in the presence of CO2 at temperatures
between (100 and 150 ºC) (Vega et al, 2014). More thermal degradation occurs as temperature
increases. At 100 ºC around 2% of MEA degrades, but 90% of MEA degrades at 150 ºC (Novitsky
et al, 2023). Thus, it is crucial to carefully control and maintain the stripper’s temperature. Thermal
degradation products mainly include OZD, HEEDA, HEIA, UREA and oligomers (Novitsky et al,
2023), these also cause corrosion of equipment and decrease MEA’s capacity and efficiency. All
thermal and oxidative degradation mechanisms of MEA are shown in section 3.2.5.
Since the products of oxidative and thermal degradation can cause equipment corrosion, it becomes
crucial to purge the degraded MEA. This can be done by removing a small fraction of the MEA
solution from the stripper unit and replace it with fresh MEA fed into the absorber. The purged MEA
can be either reclaimed or discarded, depending on the economics and environmental regulations.
The purging process can be controlled by monitoring the concentration of the degradation products
in the MEA solution. The degraded MEA can be distinguished from the normal MEA by measuring
the pH, conductivity, density, viscosity, and CO2 loading of the MEA solution. These parameters can
indicate the degree of degradation and the presence of degradation products (Braakhuis & Knuutila,
2023).
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2.2.4 Chemicals physical properties
Table 2.2: physical properties of gases and chemicals involved in Carbon Capture. (Engineering toolbox)
Component Molecular
Weight (g/mol)
Boiling
Melting
Density
at Density
at Solubility in water at
point
point
40
°C 120
°C 40 °C (g of gas/ kg of
(°C)
(°C)
(kg/m3)
(kg/m3)
H2O)
CO2
44.01
-78.55
-56.6
1.721
1.367
1
H2O
100
100
0
992.3
0.5627
-
O2
32
-183
-218.8
1.246
0.9921
0.031
N2
28.01
-195.8
-210
1.091
0.8682
0.014
SO2
64.06
-9.95
-72
2.525
1.999
55
MEA
61.08
170.3
10.3
1001
935.4
Miscible
2.2.5 Simplified Process Pathway including Main Process Units
The flue gas exits the gas turbine at 140 °C and enters a direct contact cooler (DCC) which brings
the flue gas into direct contact with cooling water to be cooled down to 40 °C. Then, the cooled flue
gas & MEA enter an absorption column at 40 °C where MEA absorbs the CO2. This loaded solvent,
saturated with CO2 is then regenerated in a stripping column. During this phase, steam (heat) at
120°C is applied to the solvent using a reboiler, which releases the captured CO 2 from the solvent,
so the solvent is returned to its original state. The lean solvent is then recycled back to the absorber,
but it is passed through a heat exchanger where heat is exchanged between the lean solvent and
the rich solvent. While the released CO2 leaves the stripper and is then cooled down and compressed
to a high pressure to reduce its volume using a train of coolers and compressors, and the CO2 is
then transported to a pipeline network which stores the captured CO2 in depleted oil/gas fields and
saline aquifers in the region. (A PBD showing all units can be seen in figure 3.1 in section 3.1).
2.3 Plant Location
It is important to consider the plant's proximity to the coal seam gas. The plant should be located
reasonably near CSG wells, and the loading centres for the generated electricity, this is to reduce
the cost of transportation and minimise the emissions of CO2 caused by transportation. However,
the plant must also be situated at a safe distance from the CSG wells in case of any unfortunate
incidents that may occur in the power station or in the CSG wells.
In Queensland, most of the coal seam gas wells are in the Surat Basin fields located in the southeast
of the state between Dalby and Miles as can be seen in figure 6. Furthermore, leading distributors
of chemicals such as Redox and Brenntag Australia which supply MEA are located in Brisbane which
is only 120 miles away from Dalby. Therefore, FAABS recommends West Dalby as the optimal
location for the proposed plant, as the area boasts close proximity to abundant coal seam gas
reserves and MEA suppliers, minimising transportation costs and logistical complexities.
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Figure 2.6: A map of Queensland showing the location of CSG wells. (Department of natural
resources & mines, 2017).
Furthermore, an assessment of Queensland’s CO2 geological storage capacity of depleted oil and
gas fields was carried out in 2008. Results showed that the depleted fields of Bowen and Surat
basins can capacitate 106 Megatons of CO2; the study also estimated the maximum potential
storage capacity of Bowen and Surat fields to be 3242 Megatons (Queensland Government, 2008).
Based on calculation in section 6.1.1, 0.367 Mt of CO2 will be captured yearly, which means that
these depleted oil and gas fields will be able accommodate our captured CO2 for the long-term.
Furthermore, West Dalby’s proximity to those depleted fields is ideal to cut down the cost of the
piping network that would be needed to transport the captured CO2.
Furthermore, West Dalby provides favourable geological conditions, ensuring stability for the
construction and operation of the power station (Mond, 1973). The existing infrastructure in the
region, including well-established roads and railways, enhances accessibility for the transportation
of materials and facilitates efficient connectivity to the electricity grid (Gov.au, 2021). Finally, due to
favourable thermodynamic conditions, most gas turbines operate at their peak efficiency when the
ambient temperature is around 15 °C (Rahman et al, 2010). This temperature optimises the
behaviour of air as an ideal gas, which enhances air density for higher mass flow rates, and improves
the efficiency of the compression process (Rahman et al, 2010). The lower ambient temperature
also supports more efficient combustion, allowing for better control of the air-fuel mixture and
combustion temperature. As can be seen from figure 7, Dalby’s average temperature is ideal for the
operation of the proposed power plant for approximately half of the year.
Figure 2.7: Climate data for Dalby taken from 1991 to 2021 (Dalby Climate, 2023).
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2.4 Environmental Impact
Carbon Capture and Storage (CCS) is a significant technology in reducing greenhouse gas
emissions from power plants, thus mitigating climate change impacts. However, its implementation
has environmental and energy efficiency implications. The CCS process requires water for chemical
and physical processes to capture and separate CO2, potentially leading to water stresses. There’s
also a risk of groundwater contamination due to potential CO2 leakage during geologic sequestration.
If CO2 escapes into water sources, it can form carbonic acid, leading to the acidification of
groundwater, the acidification process may result in the release of dissolved metals, impacting
groundwater quality and posing risks to both environmental and human health (Gustin, 2010),
especially if the groundwater serves as a drinking water source. The corrosion of underground
infrastructure, such as pipes and wells, is another consequence, potentially leading to the release of
contaminants. To mitigate these potential effects, it is imperative to implement preventive measures
during carbon capture and storage activities, including careful site selection, robust monitoring
systems, and effective well-sealing techniques. Regular monitoring and prompt detection of any
leaks are essential to enable timely remedial actions and uphold the integrity of groundwater
resources. The establishment of stringent regulatory frameworks is vital to guide the safe
implementation of carbon capture technologies and minimise risks associated with CO2 leakage into
groundwater.
On the energy efficiency front, the increase in energy demand associated with CCS can increase
operating costs. The process requires energy in the form of steam and electricity, which reduces a
power plant’s electric power output and/or increases its fuel input. This creates an energy penalty
for power plants, causing the price of electricity sold to the consumer to increase.
Finally, as mentioned in section 2.2.3, Some degradation products MEA, such as nitrosamines and
nitramines, raise health concerns due to their carcinogenic properties. Modifying process conditions
becomes crucial to prevent or minimise the production and emission of these harmful compounds.
It is crucial to control nitrosamine concentration within the range of 13.7 to 14 mM to prevent health
hazards (Badr et al., 2017). To ensure compliance with regulatory occupational exposure limits, it is
recommended to conduct exposure analyses, monitoring chemical concentrations at the plant.
Solvent losses due to oxidative degradation are significant, whereas losses from thermal
degradation are minimal. The impact of degradation on the Life Cycle Analysis (LCA) is noteworthy,
as continuous compensation for the solvent is required to maintain steady-state functionality.
It’s crucial to carefully consider these environmental and efficiency factors in the planning and
operation of carbon capture plants for power plants, aiming to balance significant reductions in
greenhouse gas emissions with minimising the environmental impact and energy costs associated
with the process.
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3. Design OF Selected Process
3.1 Process Description & Flowsheet Design
Figure 3.1: A detailed Process Block Diagram of the process route (Drawn by author).
3.1.1 Pre-treatment Units
The first step considered in the process of capturing CO2 is the pre-treatment of the flue gas. As
mentioned in section 2.2.3, SO2 can cause the degradation of MEA and form bicine. However, SO2
concentrations that causes the formation of bicine is about 10 ppm for 30 wt% MEA solution (Zhou
et al, 2013), but some studies show that SO2 at 5 ppm can actually inhibit the oxidative degradation
of MEA by O2 by scavenging oxidative radicals (Zhou et al, 2013). The removal of SOx can be
achieved by implementing desulphurisation units that utilise limestone-based scrubbers. However,
based on calculations in section 6, only 0.3 ppm of SO2 is present in the flue gas leaving the gas
turbine which means that it is very unlikely for MEA to degrade due to the presence of SO 2 in this
proposed plant. Thus, a desulphurisation unit will not be needed. Moreover, as mentioned in section
2.2.3, MEA can also degrade by the presence of NOx to form nitrosamines and nitramines. However,
MEA is a primary amine which is unable to form a stable nitrosamine, but experiments show that
under the influence of NOx, MEA degrades to the secondary amine DEA which is then nitrosated
(Fostås et al, 2011). Different techniques are used for the removal of nitramines and nitrosamines,
such as sorption through activated carbon, precursor pre-oxidation, UV irradiation, bio-treatment, &
polymerisation (Krasner et al, 2013). However, it is better to prevent/minimise the formation of
nitrosamines and nitramines by implementing a DeNOx pre-treatment unit to separate the NOx from
the flue gas before it enters the absorber. NO2 is of particular concern, it has been suggested that
concentration of NO2 in flue gas should be restricted to approximately 20 ppm (Gibbins, 2022).
Without being given the furnace conditions for the specific gas turbine that will be used by AE for the
proposed plant, FAABS cannot assess the amount of NOx that would form in the accurately.
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However, in general when combusting CSG in a gas turbine, 2-20 ppm of NO2 and 20-220 ppm of
NO form (Pavri & Moore, 2013). Therefore, FAABS believes that a DeNOx pre-treatment unit will not
be needed but recommends leaving space for a DeNOx pre-treatment retrofit in case if the NO2
emissions from the gas turbine get close to 20 ppm.
3.1.2 Direct Contact Cooler (DCC-101)
In the absence of any beneficial pre-treatment units. The flue gas exits the gas turbine at 140 °C and
1.2 bar with a H2O composition of 16.52 mol%, the flue gas is directly fed into the DCC from the
bottom, while the cooling water stream at 23 °C and 3 bar is fed into the DCC from the top as
illustrated in figure 3.1. This configuration exploits the upward movement of the hot flue gas due to
its lower density. As the flue gas rises upwards through the DCC, it comes into direct contact with
the downward-flowing cooling water. This counter current flow setup ensures a significant
temperature gradient between the ascending flue gas and the descending cooling water, promoting
effective heat transfer. The flue gas then leaves the DCC from the top at 40 °C with a H2O
composition of 15.68 mol%, the decrease in the molar composition of H2O is due to the condensation
of some water vapour from the flue gas. Furthermore, ~2 mol/s of CO2 get absorbed by the cooling
water due to Henry’s Law. The condensed water from the flue gas and the absorbed CO2 exit the
DCC along with the cooling water from the bottom of the DCC at 40 °C. This cooling water will
operate in a closed-circuit system, meaning that the water that leaves the DCC will be properly
treated then cooled using a cooling tower which is the most efficient and cost-effective cooler to use
alongside a pump to transport the water back into the DCC (Brandl et al, 2016). The cost savings
associated with decreased water supply and disposal contribute to the economic viability of the
system. Moreover, recycling also aligns with sustainable practices by reducing resource
consumption and waste generation. Finally the water that condenses must be purged in order to
prevent the accumulation of water in the system. After proper treatment this water can be used in
the power plant or externally (Magneschi et al, 2017). Finally, the DCC causes a pressure drop of
0.1 bar (Pourahmad et al, 2021), causing the flue gas to exit the DCC at 1.1 bar. Thus, the flue gas
is sent to a blower to get slightly compressed, this is to ensure that the flue gas can enter the
absorber column and flow through its packed tower structure (Mishra, 2014).
3.1.3 Absorber Column (A-101)
At this stage, the flue gas stream is ready to enter the packed bed absorption column, where the
CO2 will be captured, and the cleaned flue gas will be released to the atmosphere. In the absorber,
MEA concentration in the absorption liquid (water) is at 30 wt%, this composition is used to ensure
the specified CO2 removal of 90% and to give the optimal working conditions for MEA (Bravo et al.,
2021) (Oko et al., 2018). The cooled flue gas and the recycled lean solvent streams are both supplied
to the absorber at 40°C as it is the optimal temperature for CO2 absorption by MEA (Bravo et al.,
2021). Furthermore, the cooled flue gas stream enters the absorber column at 1.2 bar and the
pressure drop is assumed to be 0.1 bar (Tikrit University, 2023). The cooled flue gas enters at the
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bottom of the absorber, while the CO2 lean MEA stream enters from the top of the absorber column.
The CO2 lean stream containing MEA and water has a lean loading value of 0.25 mol CO2/ mol MEA
(Prasad Mishra, 2014), and a CO2 rich MEA stream leaves the absorber from the bottom as rich
amine with a loading of 0.5 mol CO2/ mol MEA (Prasad Mishra, 2014). The MEA absorbs 90% of the
CO2 in the flue gas in this process. But before the lean stream enters the absorber, it is enriched
with make-up water and periodically dosed with make-up MEA as shown in figure 3.1, this is to
resupply the system in order to counter for any vaporisation and degradation losses of water and
MEA in the stripper and absorber. As the flue gas stream rises through the absorption column the
MEA stream descends through the packed beds. The absorber is designed with packed column
beds as they provide higher contact area over the plate columns for a more efficient absorption of
CO2 by MEA (Ustadi et al., 2017). When the MEA reaches the bottom of the absorber column it is
pumped into the rich/lean heat exchanger. Meanwhile, the treated flue gas goes through a demister
pad and then into a water wash. This is to prevent any significant losses of MEA from the absorber.
A demister pad sits at the top of the absorption column. As the lean flue gas containing some
absorbed MEA rises through the column, the demister pad, composed of structured fibres or mesh,
intercepts and captures liquid droplets carried by the gas (Demisterpads.com, n.d.). Through impact
and coalescence on the pad's surfaces, larger droplets form and drain down the pad due to gravity
(Mokhatab et al, 2006). This collected liquid, which includes the absorbed MEA, is directed back into
the absorption column, ensuring that the absorbent is retained within the system for continued use.
The demister's efficiency relies on its design features, optimising liquid removal while minimising
pressure drop and preventing re-entrainment of separated droplets (Purchas et al, 2002). The flue
gas then enters a water wash column where the majority of the residual MEA that may still be present
in the flue gas is recovered to be reused. In the water wash, the flue gas is brought into contact with
water to dissolve any residual MEA, which preferentially dissolves the MEA, allowing it to be
separated from the flue gas. The recovered MEA can then be returned to the absorption process,
reducing the overall loss of MEA. The water wash reduces MEA concentrations in the cleaned flue
gas from 99ppm down to an acceptable 1.7ppm. Liu et al., 2021 (These steps taken to prevent the
release of MEA into the environment are imperative for environmental protection, regulatory
compliance, and resource conservation. Releasing MEA into the environment can potentially
contaminate water sources, soil, and air, which can harm ecosystems and wildlife. Furthermore,
resource recovery contributes to the economic efficiency of the plant. Having passed all MEA
retainment filters, the treated flue gas finally leaves absorber and at the top of the washer into the
atmosphere.
3.1.4 Heat Exchanger and Cooler (H-101)
To heat the CO2 rich MEA stream up to the desired temperature for the stripper column, a shell and
tube heat exchanger is used. The CO2 rich MEA stream leaves the absorption unit at a temperature
of 63°C and flows through a centrifugal pump that increases the streams pressure from 1.1 to 2.3
bar, and enters the tube side of the heat exchanger, here the CO2 rich MEA is heated by the CO2
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lean MEA stream and leaves the heat exchanger at 102°C this stream then flows into the stripper
unit. The CO2 lean MEA stream leaves the stripper at a temperature of 120°C, from here it enters
the shell side of the heat exchanger and is cooled to a temperature of 82.5°C. The pressure drop
across the heat exchanger on both the shell and tube side streams is assumed to be 0.5 bar
(Mukherjee, R, 1998.). Notably the pressure of the CO2 lean MEA stream is 1.7 bar, this means that
there is no required pressure increase via pump since the absorber operates at 1.2 bar. Because
the CO2 lean MEA stream is only cooled to 82.5°C, further cooling is required to cool the stream
down to the required 40°C. For this a cooler is utilised to cool the stream down as shown in figure
3.2 in section 3.1.6, after cooling the CO2 lean MEA stream re-enters the absorber.
3.1.4 Stripper Column (S-101)
In the stripper, the CO2 rich MEA stream enters the unit from the top left of the unit where it makes
its way through to the reboiler inlet at the right side of lower half of the stripper. The stripper’s
temperature is set to 120°C, this is done regenerate the solvent. The temperature is not raised further
to avoid solvent degradation and corrosion problems. In the reboiler, high amounts of steam at a
temperature of 120°C are fed to evaporate the solvent, the evaporated solvent travels into the boil
up stream and into the column where, due to heat transfer, it evaporates more of the liquid solvent
as it goes down the column. The vapour is then accumulated at the top of the stripper column where
it exits to the condenser (Oko et al., 2018). The overhead partial condenser splits the vapour into
liquid and vapour fractions and is maintained at 40°C to help the condensation process of MEA and
water. The liquid fraction, containing mostly water and MEA, is pumped back into the stripper (Oko
et al., 2018). While the vapour fractions, containing mostly CO2 and small amounts of water and
MEA leaves the system into the next step of the process. Meanwhile the lean solvent is pumped
back into the heat exchanger to start the cycle again. It should be noted that the average pressure
in stripper and the condenser should be 1.8-1.9 bar, while the pressure in the reboiler is 2 bar. This
is done to ensure best performance out of the unit and the highest rate of solvent regeneration (Bravo
et al., 2021) (Prasad Mishra, 2014). The loading of the inlet lean amine is around 0.25 and its leaves
with CO2 as rich amine with a loading of 0.5, the difference between these two values is the amount
of CO2 regenerated in the stripper. Furthermore, the pressure drop in the stripped is assumed to be
at the average of 0.1 bar (Prasad Mishra, 2014) (García et al, 2022). Notably, the most energy
intensive part of the process is the solvent regeneration in the reboiler.
3.1.5 Compressor Train (COM-101 TO C-105)
After the separated CO2 leaves the stripper, it heads into the last step of the process which is to be
cooled and compressed to 25 °C and 50 bar in preparation to be put in storage. However, due to
some water being present in the stream, the water content in the separated CO2 flow must first be
removed through the use of a dryer. This is because water and CO2 are corrosive when mixed
(Onyebuchi et al., 2018) . After the water has been removed from the stream, the concentrated CO2
enters the first of four compressors at 1.7 bar and 40°C. Four compressors are used instead of three
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to lower the interstage pressure between each compressor, which decreases the temperature rise
that the stream experiences and lowers the individual duty of each compressor and cooler, at the
cost of an extra compressor and cooler. The stream is then compressed to 14 bar and experiences
a sharp rise in temperature, so it is then sent to the first cooler to decrease the temperature. Cooling
water at 23°C absorbs heat from the CO2 stream, due to a temperature difference. This process is
repeated three more times, with each compressor increasing the pressure of the stream by 12 bar,
until the stream is at 50 bar. The stream finally enters the fourth cooler, where it is cooled to the
desired temperature of 25°C where it is then ready to be sent to the pipeline for storage.
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3.1.6 Process Flow Diagram
Figure 3.2: PFD of the process with streams conditions
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3.2 Process Chemistry & Thermodynamics
3.2.1 Absorption of CO2 by water
As mentioned in section 3.1.2, some CO2 is absorbed by the cooling water used in the DCC to cool
down the flue gas. The process of CO2 absorption by water involves a series of chemical reactions
and is influenced by thermodynamic principles. When CO2 comes into contact with H2O, it can
dissolve to form carbonic acid (H2CO3). This reaction occurs when the pH is in the range of 3-8, but
the reaction has slow kinetics (Britannica, 2023). Moreover, this is a reversible reaction, meaning
the carbonic acid can decompose back into water and carbon dioxide. The reaction can be
represented as shown below:
CO2 + H2O ⇌ H2CO3 (3.1)
Furthermore, carbonic acid is a weak acid and can dissociate in water to form bicarbonate (HCO3-)
and a hydrogen ion (H+) (Britannica, 2023). The HCO3- is more stable than and less likely to release
the CO2 back into the gas phase (Britannica, 2023). Furthermore this reaction occurs when the pH
is in the range of 3-8, and the reaction has fast kinetics (Britannica, 2023). This reaction can be
represented as shown below:
H2CO3 ⇌ HCO3− + H+ (3.2)
Moreover, Bicarbonate can further dissociate to form carbonate ions (CO32-) and another hydrogen
ion (Britannica, 2023). This reactions has fast kinetics but this reaction is not likely to occur as it
requires a pH higher than 10. This reaction can be represented as shown below:
HCO3− ⇌ CO32−+H+ (3.3)
These reactions are equilibrium reactions, meaning they can proceed in both directions. The position
of the equilibrium (how much of the reactants convert into products) is influenced by factors such as
temperature, pressure, and the concentration of the reactants and products. In terms of
thermodynamics, the absorption of CO2 in water is an exothermic process, meaning it releases heat.
The solubility of gases in water typically decreases with increasing temperature, so more CO2 will
dissolve in cold water compared to warm water. Furthermore, the absorption of CO2 by water is
significantly influenced by pressure. According to Henry’s Law, the solubility of a gas in a liquid is
directly proportional to the partial pressure of the gas above the liquid. This means that as the partial
pressure of the CO2 increases, more CO2 will dissolve in the water. Henry’s law was used in section
6 to calculate the amount of CO2 that gets absorbed by the cooling water stream:
𝐶 = 𝑘𝐻 × 𝑃𝐶𝑂2 (3.4)
Henry’s constant (kH) was taken at 23 °C (0.034 mol/atm) from literature (NIST, 2023). This is
because Henry's law is typically applied to the temperature of the liquid phase, which, in this case,
would be the water at 23 °C.
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Furthermore, the O2 and N2 present in the flue gas can also be absorbed by the cooling water.
However, as can be seen from table 2 in section 2.3.1, the solubility of O 2 and N2 are very low
compared to CO2. Thus, their absorption by water is neglected.
3.2.2 Absorption of CO2 by MEA Kinetics & Mechanisms
The main reaction in the absorption column involves absorption of CO2 by the aqueous MEA. As
mentioned in section 2.2.2, the carbon atom of CO2 binds with the nitrogen atom of MEA via a
nucleophilic attack, forming a Zwitterionic intermediate (MEA+COO-) (Holmes, 2018). The reaction
can be represented as follows:
k2
C2H4OHNH2 + CO2 ⇌ C2H4OHNH2+COO- (3.5)
k-1
Furthermore, the Zwitterionic intermediate then gets deprotonated by water molecules, forming a
stable carbamate ion and a solvated proton. The reaction can be represented as follows:
𝑘𝑏
C2H4OHNH2+COO- + H2O → C2H4OHNHCOO- + H3O+ (3.6)
Then another MEA reacts with the solvated proton as mentioned in section 2.2.2. The mechanisms
of these reaction are all shown in figure 3.3.
Figure 3.3:
Mechanisms involved during CO2 capture in aqueous MEA (Hwang et al.,2015).
The rate of reaction can be expressed as shown below:
−rCO2, C2H4OHNH2 =
k2 [CO2 ][C2 H4 OHNH2 ]
(1+∑
k−1
)
kB [H2 O]
(3.7)
For the CO2 absorption by MEA system, the overall reaction rate is second order, being first order
with respect to both CO2 and MEA (Sakwattanapong et al., 2009). Therefore, equation 3.7 can be
rewritten as show below:
−rCO2 , C2 H4 OHNH2 = k 2 [CO2 ][C2 H4 OHNH2 ] (3.8)
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Finally, based on experiments, researchers determined a value for k2. According to Hikita et al.
(1977), the kinetic data can be characterised by the equation shown below:
log(k 2 ) = 10.99 −
2152
T
(3.9)
3.2.3 Whitman Two-Film Mass Transfer Theory
In an absorber packed column in a post-combustion CO2 capture plant, the efficiency of CO2
absorption by MEA is influenced by factors such as the gas-liquid contact degree, physicochemical
properties, hydrodynamics of the absorber column, amine reactivity degree, and operating
parameters related to the gas and amine solution (Khan et al., 2011). The chemical absorption of
CO2 into an amine solution is often described by Whitman’s two-film theory, which proposes the
existence of two thin films near the gas and liquid phase interfaces. These films, separating the
phases from their respective bulk phases, assume that the bulk phases are in equilibrium and that
all mass and heat transfer resistances exist within the two films (Ghadiri, 2023). When CO2 moves
from the gas to the liquid phase, a chemical reaction between CO2 and the amine solution can take
place in the liquid film or liquid bulk, fast reactions occur in a narrow zone within the film, while
slow reactions spread through the film and liquid bulk (Afkhamipour, 2017). This is dependent on the
relative values of the reaction rate constants, mass transfer coefficients of gas and liquid phases,
concentration ratio of reactants and CO2 equilibrium solubility (Afkhamipour, 2017).
3.4: schematic diagram for Whitman’s two-film model, showing the location of the chemical
reaction between CO2 and the amine (Afkhamipour, 2017)
Figure
Furthermore, the reaction between CO2 and MEA is assumed to be infinitely fast, the reaction is fast
enough so that this reaction zone remains totally within the liquid film (Levenspiel, 1999). Thus, no
CO2 enters the main body of liquid amine to react there. Therefore, the overall rate of absorption
of CO2 in an aqueous solution of MEA considering the mass transfer and chemical reaction rates
can be expressed based on the two-film model for a fast second-order reaction as shown below:
−rCO2 =
pCO2
(k
HCO
2 )
+
CO2 g 𝑎 kCO2 l a E
1
(3.10) (Levenspiel, 1999)
In the CO2 absorption system, Henry’s constant is very large (Ghadiri, 2023), which means that:
1
k CO2 g 𝑎
≪
HCO2
k CO2 l a E
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Therefore, it could be considered that the resistance to mass transfer resides in the liquid film,
neglecting the mass transfer resistance by the gas film (Ghadiri, 2023).
3.2.4 Vapor-Liquid Equilibrium (VLE) Data
3.2.4.1 Optimum temperature for absorption
VLE Data is a fundamental aspect of understanding the behaviour of CO2 in MEA solutions. This
data provides insights into the phase behaviour of the CO2-MEA-H2O system under various
conditions of temperature and pressure. It is crucial for the accurate simulation and design of CO2
absorption processes; it allows for the prediction of the system’s performance under different
operating conditions.
3.5: CO2 partial pressure as a function of loading for 30 wt% MEA at different temperatures.
(○=40 °C), (◊=60 °C), (□=80 °C), (Δ=100°C), (∇=120°C) (Aronu et al.,2011).
Figure
Figure 3.5 shows that, at a temperature of 40 °C, the absorption of CO2 by 30 wt% MEA exhibits the
lowest CO2 partial pressure requirement to attain a specific CO2 loading value at any given point.
This suggests that 40 °C is the optimum temperature for CO2 absorption by 30 wt% MEA. This is
because when a lower CO2 partial pressure is required for the same loading, it suggests a
heightened sensitivity to capture CO2 from flue gas streams which is beneficial for streams with a
low CO2 concentration (mol%), facilitating more effective capture (Khan et al., 2023). Moreover, a
lower CO2 partial pressure requirement means that the absorber column requires a lower total
pressure (Ravi, 2023). This reduces the need for compressing gases to high pressures, leading to
energy savings and decreasing operational costs. Furthermore, it also allows for the use of less
expensive and more readily available equipment, as well as materials with lower pressure ratings
(King, 2020).
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3.2.4.2 Optimum pressure for absorption
Table 3.1: VLE date from literature for 30 wt % aqueous MEA at 40 °C (Jayarathna et al., 2013).
This table can be further illustrated as a graph as shown in figure 3.6:
Figure 3.6: A graph of the VLE data shown in table 3.1.
As shown in table 3.1 and figure 3.26, a higher CO2 loading requires a higher partial pressure of
CO2. Increasing the partial pressure of CO2 approximately from 0.01 to 10 kPa increases the CO2
loading approximately from 0.2 to 0.5 mol CO2/mol MEA. However, a further small increase of the
CO2 loading from 0.49 to 0.58 requires a huge increase in CO2 partial pressure from 10.75 to 82.7
kPa (Lee et al., 1974). This is mainly because as mentioned on section 2.2.2, two moles of MEA are
needed to absorb one mole of CO2 theoretically, meaning that the loading value is 0.5 mol CO2/mol
MEA. However, with enough high partial pressure, a strong driving for more CO2 absorption by MEA
can be created, allowing for the loading value to exceed 0.5 mol CO2/mol MEA (Zhang et al., 2017).
While increasing the partial pressure of CO2 to high pressures can enhance its loading value in MEA
absorption, industries typically avoid doing so due to several reasons. Firstly, raising the CO2 partial
pressure significantly increases energy consumption, leading to higher operational costs. Secondly,
equipment designed to withstand high pressures tend to be more robust and expensive, increasing
the capital costs (King, 2020). Lastly, safety concerns also arise with high pressures, including the
20
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risk of equipment failure or leaks that could lead to the release of CO2 and other gases (Cusco,
2007). Based on the VLE data, the optimal partial pressure of CO2 is identified at 10 kPa. This choice
is a compromise, maximising absorption efficiency while managing the associated costs, considering
factors such as energy requirements, operational costs and expensive robust equipment
requirements. Based on calculations in section 6, the cooled flue gas entering the absorber column
has a composition of 0.0855 mol%. The optimal total pressure inside the absorber column is
calculated using the equation below:
𝑃𝐶𝑂2
𝑃𝑇𝑜𝑡𝑎𝑙 = 𝑋
𝐶𝑂2
𝐴𝑏𝑠𝑜𝑟𝑏𝑒𝑟 𝑂𝑝𝑡𝑖𝑚𝑢𝑚 𝑃𝑇𝑜𝑡𝑎𝑙 =
(3.11)
10
≈ 120 𝑘𝑃𝑎 = 1.2 𝑏𝑎𝑟
0.0855
3.2.4.3 Optimum MEA wt%
Figure 3.7.
A graph of the VLE data at 40 °C for MEA with different concentrations (wt%) (Aronu et
al.,2011).
As shown in figure 3.7, the absorption of CO2 by very low wt% MEA exhibits the lowest CO2 partial
pressure requirement to attain a specific CO2 loading value at any given point. However, using a
very low wt% of MEA causes a significant need to increase the solvent flow rate (Wang et al., 2023).
This is because despite the potential for higher CO2 loading values with lower MEA concentrations,
each unit of solvent has a diminished capacity for CO2 absorption (Ravi, 2023). Therefore, this
limitation necessitates a larger volume of solvent circulation to capture equivalent amounts of CO 2.
Higher flow rate requirements in industrial processes lead to increased costs due to the need for
larger equipment, elevated energy consumption for pumping, higher infrastructure and maintenance
expenses, and potential compliance-related expenditures (King, 2020). Furthermore, the mass
transfer rate also decreases at lower MEA concentrations (Wu et al., 2017), which necessitates a
greater time of contact between the solvent and the flue gas, which means that an absorber column
with a greater column height would be required, increasing the capital costs. Additionally, the
regeneration of the solvent for reuse after CO2 capture often requires more extensive processing at
lower MEA concentrations, this is due to the higher flow rate used with low wt% MEA (Wang et al.,
2023). This higher flow rate requirement not only impacts energy consumption but also poses
operational and economic considerations.
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On the other hand, using a very high wt% of MEA in CO2 absorption processes offers a higher CO2
absorption capacity per unit of solvent (Akram et al., 2020), resulting in increased efficiency and
potential reductions in solvent flow rates and operational costs. However, challenges arise as higher
MEA concentrations may lead to increased viscosity (Amundsen et al., 2009), impacting fluid
dynamics within the absorption column and requiring more pumping power, increasing the costs for
pumping. Moreover, high MEA concentrations can elevate the corrosivity of the solvent (Akram et
al., 2020), necessitating corrosion control measures and the use of corrosion-resistant materials
which increases the costs. Furthermore, the cost of MEA is a significant factor in the overall
economics of the absorption process and using higher a wt% of MEA directly increases the costs of
the absorption process. A 30 wt% of MEA in CO2 absorption is considered optimal as it is a
compromise between optimising absorption performance within the constraints of operational and
economic considerations (Zhang et al, 2020).
3.2.5 Side Reactions Kinetics and Mechanisms
3.2.5.1 Oxidative degradation of MEA by SO2
As mentioned in section 3.1.1, oxidative degradation of MEA by SO2 can occur to form bicine.
However, this requires at least 10 ppm of SO2 for 30 wt% MEA solution to be present in the flue gas
(Zhou et al, 2013). As shown in section 6, the proposed plant only has 0.3 ppm present in the flue
gas, making the formation of bicine highly unlikely. The rate of reaction for the degradation of MEA
by SO2 can be expressed as shown below (Supap et al., 2008):
𝐸𝑎
−rC2H4 OHNH2 =
−
k0 𝑒 𝑅𝑇 [C2 H4 OHNH2 ]𝑎 [O2 ]𝑏 [SO2 ]𝑐
1 + kB [CO2 ]𝑑
(3.12)
Based on experimental results obtained by Uyanga et al. (2007), equation 3.12 can be rewritten as
shown below:
20752
−rC2H4 OHNH2 =
−
0.7189𝑒 8.314𝑇 [C2 H4 OHNH2 ]1.359 [O2 ]0.03 [SO2 ]2
[CO2 ]0.033
(3.13)
3.2.5.2 Oxidative degradation of MEA by NOx
Furthermore, as mentioned in section 3.1.1, photo-oxidative degradation of MEA by NOx can occur
to form nitrosamines and nitramines when NO2 concentrations are around 20 ppm in the flue gas.
However, MEA is a primary amine which is unable to form a stable nitrosamine, but under the
influence of NOx, MEA degrades to the secondary amine DEA which is then nitrosated (Fostås et al,
2011). The mechanisms of the MEA degradation by NOx are shown in figure 3.8.
3.2.5.3 Oxidative degradation of MEA by O2
Moreover, as mentioned in section 2.2.3, oxidative degradation by O2 can also occur in the absorber
if the flue gas O2 composition is around 3% or higher (Vega et al, 2014). The main products of MEA
oxidative degradation are ammonia, formic acid, acetic acid, HEF, HEI, and HEPO. The different
MEA degradation reactions are shown in figure 3.9.
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Figure 3.8. MEA degradation by NOx mechanisms (Vevelstad et al, 2011)
Figure 3.9. MEA oxidative degradation
reactions (Vevelstad et al, 2011).
3.2.5.3 Thermal degradation of MEA
Finally, thermal degradation also inevitably occurs in the stripper in the presence of CO2 at
temperatures between (100 and 150 ºC) (Vega et al, 2014). More thermal degradation occurs as
temperature increases. At 100 ºC around 2% of MEA degrades, but 90% of MEA degrades at 150
ºC (Novitsky et al, 2023). Thermal degradation products mainly include OZD, HEEDA, HEIA, UREA
and oligomers (Novitsky et al, 2023). The mechanisms of thermal degradation reactions are shown
in figure 3.10.
Figure 3.10.
MEA thermal degradation pathways (Jens, 2013).
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4. RAW MATERIALS AND PRODUCTS
4.1 Raw Materials
4.1.1 Raw Materials for Construction:
Table B.1 in Appendix B provides a list of international and local steel and concrete suppliers for the
construction of the carbon capture plant. Both international and local suppliers are considered as
they both have their separate advantages and disadvantages. International steel and concrete
suppliers should be used as the company can benefit economically from the reduced prices and
advanced technology they offer, due to them being much larger corporations. Both steel and
concrete should be supplied from China. From China, Baowu Steel Group and China National
Building Material Group respectively.
Table 4.1: Advantages and disadvantages of local and international suppliers
Local suppliers
Advantages:
Disadvantages:
-
-
Increased flexibility and decreased leadtime when purchasing raw materials from
a local supplier.
It is likely going to be an easier process
dealing with a local supplier compared to
an international supplier (e.g. language
barriers, time-zone differences)
International suppliers
Advantages:
-
-
-
Likely to be smaller suppliers that can’t
match international supplier’s prices or
quality.
Specific raw materials may not be
available from a local supplier.
Disadvantages:
-
Can offer lower prices and better-quality
products due to the company being larger
in size.
There are a greater number of international
suppliers compared to local suppliers that
can meet the specific requirements for the
raw materials.
-
Increased transportation cost of raw
materials.
Political and economic issues in the
country or region of the supplier can impact
the supply of raw materials.
4.1.2 Raw Materials for Operation:
Referring to Figure B.1 in Appendix B, it is illuminated that the price of MEA is expected to decrease
over the coming years. This means that it is likely economically beneficial to not enter a long-term
contract with any MEA suppliers and rather consider signing yearlong contracts to leverage the
decreasing market price of MEA. Furthermore, the market price of MEA in SE Asia is currently and
is projected to be the lowest out of all other areas so if an international supplier is to be chosen, it
should be from the SE Asia region. Furthermore, this would decrease the transportation costs of the
MEA from the supplier to the carbon capture plants due to the geographical location of SE Asia.
However, due to current political factors in the SE Asia region, such as border disputes and conflicts
as well as high levels of corruption, it may be more beneficial in the long-term to opt for a supplier in
a more politically stable region (Heng, 2020). A suitable choice is SABIC, an international MEA
supplier based in the Middle East, which has the second lowest MEA market price according to figure
appendix B.3.
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As the price for MEA is similar for both local and international suppliers, it is more economically
advantageous to opt for a local supplier due to the decreased transport costs and for ease of
communication. A suitable local MEA supplier would be Redox Pty Ltd., which operates a short
distance away from the carbon capture plant. This would also allow for more flexibility when it comes
to acquiring the MEA and a shorter lead-time, as well as the potential for a good relationship with
Redox Pty Ltd. as a long-term MEA supplier.
4.1.3 Coal bed methane supplier:
A local supplier for coal bed methane could be beneficial as this avoids any global events that might
affect the supply chain, such as COVID-19, which impacted global trade. Due to the creation of new
policies and laws during the lockdown period, this significantly affected shipping networks and ports
which resulted in challenges regarding international trade (World Bank, 2022). Similarly, events such
the Russia-Ukraine war, have affected the European and Australian imported fuel market heavily.
Australia has long imported gas from Russia in the past due to the cheap prices offered however on
the 25th of April 2022, the Ministry of Foreign Affairs in Australia issued a ban on the acquisition and
transportation of Russian-based oil, gas, petroleum derivatives and coal (Department of Foreign
Affairs and Trade, 2023). As global political climates are constantly changing, it could be beneficial
to acquire coal bed methane from a local supplier, such as Woodside Energy which is one of the
largest oil and gas suppliers in Australia, making them a low risk and potential long-term supplier
(Bellantuono et al, 2023).
Although, it would be very beneficial to have a local supplier for coal bed methane economically, this
is not feasible. Due to an initiative from the Australian fracking industry to build new fracking fields,
it costs between 3.65 - 6.40 Australian dollars to produce one gigajoule of energy from coal seam
gas which is much higher than the expected price of 2.20 - 2.70 Australian dollars per gigajoule. As
gas price is an essential player in the Australian electricity market, electricity prices in Australia,
specifically in Queensland where the plant is located, have skyrocketed after the fracking initiative
with domestic gas prices rising to a peak of 21 Australian dollars per gigajoule in 2017 (Robertson,
2019). It is much more economically favourable to have a coal bed methane supplier from China
where it costs 0.2 American dollars per metre cubed (Yihe, 2021). All things considered; it is best to
use China United coalbed Methane Corp. Ltd. as the coal bed methane supplier. For coal bed
methane, it is best to obtain a long-term contract as the price of coal bed methane is currently
increasing as the coal bed methane market is emerging, with prices likely to rise as the global coal
bed methane market expecting to grow by 4.9% during 2023-2028, from US$ 18.9 Billion in 2022 to
US$ 25.4 Billion by 2028 (imarcgroup, 2023).
In addition to this, new legislations or policy changes could lower the product prices, such as
Australia aiming to have a zero net emission of carbon by 2050 could lead to an increase in the price
of coal bed methane due to a potential increase in carbon tax in Australia, which would see
organisations moving away from traditional oil as a fuel.
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4.2 Product and By-products
After capturing of the carbon dioxide, it is then transported through pipelines to be stored
underground. The utilisation of the captured carbon dioxide is referred to as CCUS (carbon capture,
utilisation, and storage), which has a global market value of $1.9 billion in 2020 and is projected to
reach $7 billion by 2030 (alliedmarketresearch, 2020). CO2 is used in industries such as the food &
beverages, fire extinguishers, refrigeration and polymers industries to be sold as a product (Co,
2021). The food & beverage industry is stable and is projected to be worth $250 billion in 2030,
making them a large consumer of CO2 (CSIRO, 2019). Diverting captured CO2 for commercial use
in the food and beverage sector may contribute to increased competition for a finite resource,
possibly impacting other crucial industrial applications and exacerbating market dynamics (Mackey
et al, 2013). One of the major applications of CCUS is enhancing oil recovery which is done by
injecting carbon dioxide into oil reservoirs to increase the amount of oil that is capable of being
extracted. However, most of the injected carbon dioxide enters the atmosphere once again
(McGlade, 2019). The captured carbon dioxide can also be converted into cement through the
process of mineralisation (Ralston, 2021). This does not release the carbon dioxide back into the
atmosphere. However, not much carbon dioxide will be needed for this process. The price of CO2
varies on its purity; the captured CO2 would have a market price of 85$ per ton or 60$ per ton if used
in enhancing oil recovery (MIT, 2023).
Referring to figure B.2 in Appendix B, the demand for carbon dioxide is likely to increase in the
coming decades. This is due to increased efforts in enhancing oil recovery and the very stable
agriculture sector, specifically the fertiliser industry, which requires carbon dioxide to produce urea .
However, in 2021, the amount of carbon dioxide captured globally was estimated to be around 40
million tonnes per year, across 26 operational CCS facilities (Brogan, 2022). Moreover, the
increased efforts in implementing CCUS units should increase the amount of carbon dioxide
captured in the coming decades. To maximise use of the captured carbon, it should be stored to
obtain carbon credits as Australia has a government-run carbon credit scheme which gives
Australian Carbon Credit Units (ACCUs). One ACCU is earned for each tonne of CO 2 stored by a
project. This project is projected to capture and store 370,000 tons of CO2 per year, worth more than
$4.5 million (Starr, 2023). Furthermore, this approach aligns with the imperative of achieving climate
targets. As there are emerging applications for carbon dioxide, the CCUS market is competitive with
many corporations starting to build units to meet the increasing demand for carbon dioxide. This
should increase efficiency and advance technological operations to maximise the amount of carbon
captured (MIT, 2023).
5. SAFETY, HEALTH, AND ENVIRONMENT
5.1 Process Safety
In order to guarantee the safety and integrity of the process, a HAZOP0/1 was carried on the
proposed carbon capture plant. The study involves explosion, corrosion, fire, extreme weather,
conventional safety and loss of service conditions and their possible consequences.
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Table 1. HAZOP0/1 Study
Guidewords
Hazard
Consequences
of HAZARD
Pipe blockage
Buildup of fluid,
high pressure,
potential
equipment
damage
upstream
Injury to
workers, release
of flue gas
stream to
atmosphere
Pump increases
pressure of
stream
excessively,
damage to
downstream
equipment
Potential
unwanted
exothermic side
reactions
Massive
damage to plant,
loss of life or
serious injury,
release of
harmful
components to
atmosphere
Reaction rate
and heat release
increased. May
eventually
exceed vessel
cooling capacity.
leading to overtemperature
Injury to
workers, release
of flue gas
stream to
atmosphere
Contamination
of stored product
Pipe rupture
Pump failure
Explosion/overpressure
Flue gas
impurities
Vessel failure
Increased flow
rate.
Pipe rupture
Corrosion
Storage tanks
Fire
Electrical
spark
Damage to
equipment.
Human error
Potential to start
fire.
27
Hazard
Management
strategy
Pressure
indicators
Actions
Comments
Consider
operating
conditions.
Regular
maintenance.
&
ESD
shutdowns on
process to be
installed
Design of
control systems
for shutdown
On-line flue gas
monitoring
Constant
monitoring of
process
conditions.
Independent
alarms
Perform regular
maintenance.
Consider
operating
conditions,
Perform regular
maintenance.
Lining the tank
with corrosion resistant
material
Ground all
pumps and
compressors
Designated
smoking areas
Perform regular
maintenance.
Activities with
potential to
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cause fire must
not be done on
plant premises.
Flammable
products. i.e.
MEA
Combustion of
hot vapours
Lightning
Power loss,
damage to
equipment, fire
Flooding
A hold back on
normal personal
operations,
Possible
damage to
equipment.
Noise pollution,
potential hearing
loss of workers
Extreme weather
Noise
Conventional safety
Electricity
Steam
Loss of service
Electricity
Risk of
electrocution,
electrical spark
causes fire
Risk of
condensation
leading to
overflow in
stripper column
and reboiler
outlet.
Interruption of
process
monitoring
causing to loss
of control over
plant.
Remove all
sources of
ignition from
areas where
flammable
products are
stored
Lightning rod
on tall
equipment,
ground all
electrical
equipment
Consider
alleviating
equipment
susceptible to
water damage.
Place pumps
and coolers
away from
workstations
Ground
equipment,
regular
maintenance
ESD of process
until issue is
resolved.
Consider
having an
electrician.
Backup power
generators.
5.2 Health
The carbon capture plant involves the handling of many dangerous materials if not handled properly.
These materials pose risks to human health which must be understood to ensure the safety of
operators, employees, and the surrounding public. In order to understand these risks material safety
data sheets were utilised to present relevant hazards to the plant proposed. The key elements of the
process will be considered in this section, these are deemed to be CO2, MEA, H2O, O2, and N2.
At the proposed plant’s conditions, CO2, is a colourless gas. At low concentrations, the gas is
odourless. At higher concentrations it has a sharp, acidic odour. As such, the handling of CO2 must
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be done with the utmost care to prevent hazards and risks. In the case of the outlet stream of the
plant, when the CO2 gas is compressed & liquefied, it poses no toxicity risk, however if pipes carrying
the component are amply heated, the risk of explosion becomes apparent, possibly endangering the
lives of close-by personal. In case of high concentrations care should be taken as CO2 is considered
an Asphyxiant.
With regards to MEA, it is the most harmful material to human life in the plant. Contact with the skin
causes severe burns, irritation to the eyes could cause serious eye injury, and exposure is likely to
cause respiratory irritation. If not handled correctly, MEA can cause serious damage to personal or
employees. Furthermore, MEA is susceptible to combustion at high temperatures (2).
O2 and N2 at this plant are do not pose any toxicological or ecological threats, however, both
components are considered and Asphyxiants. Oxygen might intensify fires, both nitrogen and
oxygen may explode if heated (Air Liquide, 2021). Water has no adverse health effects, no
toxicological or ecological effects if handled properly.
In order to fully understand the possible health effects and the way of mitigating the chances of such
risks from occurring, a materials safety data sheet was created for the proposed plant. And so, further
details around the process and all considered components including prober storage and disposal
conditions alongside accidental release measures and all other relevant information can be seen in
Table (5.2) Below. All information was taken from individual component data sheets (Air liquid, 2022.
Honeywell. 2017 CHEM supply Australia. 2020. LAB CHEM, 2020).
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Table 5.2: Material safety data sheet for all key components involved in the process
COMPONENT
CO2
MEA
H2O
O2
N2
IDENTIFIER
Carbon dioxide
CAS-No.: 124-38-9
MONOETHANOLAMINE
CAS-No.: 7727-37-9
Water
CAS-No.: 7732-18-5
Oxygen
CAS-No.: 7782-44-7
Nitrogen
CAS-No:7727-37-9
N.A.
N.A.
N.A.
None
May cause or intensify fire,
oxidiser.
May explode if heated.
May explode if heated.
No adverse health effects
No adverse health effects.
considered Asphyxiant
No toxicological or ecological
damage caused by this product.
considered Asphyxiant
Material can create slippery
conditions.
Use non-slip safety shoes in areas
where spills or leaks can occur.
Ventilate area.
In case of fire: Stop leak if safe to
do so.
Prevent from entering sewers,
basements and workpits, or any
place where its accumulation can
be dangerous.
Remove victim to uncontaminated
area.
Skin and eye contact have no
Adverse effects expected.
Adverse effects not expected from
this product
General room ventilation
No special precautions are
normally required when handling
this product
Avoid inhalation. Use in wellventilated areas. Where an
inhalation risk exists, mechanical
extraction ventilation is
recommended. Wear PPE.
Permitted disposal of water. No
special protective measures
required
Ensure no corrosive conditions.
Keep container below 50°C in a
well-ventilated place. Segregate
from flammable gases and other
flammable materials in store.
Protect from sunlight. Store in a
well-ventilated place.
May be vented to atmosphere in a
well ventilated place. Do not
discharge into any place where its
accumulation could be dangerous.
COSHH PICTOGRAM
N.A.
STEL EXPOSURE LIMIT
30,000ppm, 54000 𝑚 𝑔3
HAZARDS
May displace oxygen and cause
rapid suffocation.
May increase respiration and
heart rate.
HEALTH EFFECTS
ACCIDENTAL RELEASE
MEASURES
FIRST AID MEASURES
𝑚
No known significant effects or
critical hazards.. considered
Asphyxiant.
Evacuate. Avoid breathing gas.
Provide adequate ventilation.
Wear appropriate respirator when
ventilation is inadequate. Put
PPE.
Remove victim to uncontaminated
area. Wash skin and eyes if in
contact.
15
𝑚𝑔
, 6 𝑝𝑝𝑚
𝑚3
Harmful in contact with skin.
Causes severe skin burns and eye
damage.
Harmful if inhaled.
May cause irritation to nose, throat
and lungs in high concentrations. May
cause redness, pain and blistering.
Danger of skin absorption can cause
severe eye irritation and burns with
redness, pain.
and blurred vision.
Use personal protective equipment.
Wear respiratory protection.
Ensure adequate ventilation
If breathed in Take individual outside
into fresh air.
In case of skin contact:
remove the infected shoes and
clothing right away.
in case of eye contact:
Rinse well.
EXPSOURE CONTORLS
Good general ventilation should
be sufficient to control worker
exposure to airborne
contaminants.
Effective exhaust ventilation system
Ensure that eyewash stations and
safety showers are close to the
workstation location.
STORAGE AND DISPOSAL
Protect from sunlight. Store in a
well-ventilated place.
Dispose of surplus and nonrecyclable products
via a licensed waste disposal
contractor. Waste should not be
disposed of untreated to the
sewer unless fully compliant with
the requirements of all authorities
with jurisdiction
The product should not be allowed to
enter drains, water courses or the soil.
No smoking.
Keep container tightly closed in a dry
and well-ventilated place. Reacts with
copper, aluminium, zinc, and their
alloys
30
No hazards which require special
first aid measures.
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5.3 Environmental Impact
Globally, over 40% of energy-related carbon dioxide emissions are due to the burning of fossil fuels
for electricity generation, in Australia, the situation is worse where, in 2021, there were four tonnes
of carbon dioxide released for every Australian (World Nuclear Association, 2022). The average
person in the world only produces 1.1 tonnes of carbon dioxide per year to generate electricity
(Broadbent, 2023). The ability to retrofit a plant with post-combustion carbon capture, is forecasted
to greatly reduce the carbon emissions in Australia (Geoscience Australia, 2023). However, it is
important to note that carbon capture is not completely free of the risk of emissions.
Under normal circumstances, it is important to continuously monitor the process, the primary goal of
process monitoring is to demonstrate that the process's emissions, mostly to the air, do not pose a
threat to the environment (UK EA, 2022). In normal operations adequate safety measures and
containment protocols must be in place to prevent accidental spills or releases, all units were
optimised to minimise MEA release or degradation (see chapter 3?). However, in the case of leaks
in the process, whether they be CO2 and MEA related or not, they need to be identified and rectified.
Notably, in process dealing with similar amounts of carbon, fugitive emissions from the natural gas
industry (which encompasses both conventional and CSG gas) were estimated to be 9.3 million
tonnes of carbon dioxide in 2008–09 (DCCEEW, 2012).
In the case of accidents or failure of the operation, the hazards have a far more immediate impact.
Firstly, in the case of accidental release of the MEA can pose risks to the environment. For example.
For example, harming the aquatic life if in contact with any water streams, in cases such as this,
these spillages can be contained and then collected with non-combustible absorbent material such
as sand, earth, diatomaceous earth and vermiculite (Air liquid, 2022). The most pronounced issue
is the rerelease of CO2 into the atmosphere, this could occur in a number of ways. Firstly, if CO2 was
released in significant quantities at the plant site or in the surrounding area. Secondly, if there are
issues with underground storage, it could lead to immense environmental concerns that would affect
the surrounding water and ecosystem, agricultural land, the rerelease of great amounts of CO 2 into
the atmosphere (St John, 2013).
In terms of transporting materials, CSG deposits in Australia are frequently found near areas of prime
agricultural land, such as Queensland's Darling Downs and New South Wales Liverpool Plains (St
John, 2013). According to academic findings and media reports, there may occasionally be a conflict
between agricultural practices, the development of CSG, and the movement of martials into and out
of plants (St John, 2013). Naturally, the carbon capture plant is located within premises close to the
CSG plants where materials are typically transported by truck or rail, which can contribute to
damaging and polluting the nearby agriculture land. Furthermore, in the transportation CO2 form the
carbon capture plant to the storage site, the most unwanted impurity in the stream is considered to
be water. If not properly removed, it causes hydrate formation in the transport pipeline. Therefore, in
a similar transportation conditions used in this plant, the presence of water is likely to lead to
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corrosion and leaks or even raptures in the pipes, realising CO2 into the atmosphere (Onyebuchi et
al., 2018).
Furthermore, the process involves wastewater containing contaminants that must be treated before
discharge to prevent negative impacts on aquatic surroundings (St John, 2013). Damage through
release of untreated production water could also lead to damage to wildlife habitat in sensitive areas
and contamination of surface water resources in drinking water catchment (St John, 2013). Modelling
undertaken by the former Queensland Water Commission suggested that CSG, activities might
affect a small proportion of agricultural aquifers in the Surat basin if contaminates are not disposed
of correctly if wastewater is not disposed of properly, it would only further disrupt the local aquifers
and environment (St John, 2013). Therefore, wastewater in the plant must be disposed of properly,
one method of doing this is treating the purge water using water treatment units before purging or
by hiring waste contractors. Notably, all other by-products of this specific flue gas coming out of the
absorber are clean and within safe limits, these by-products have no effect on the personal or the
environment.
In the proposed carbon capture plant, there is 1.44 Kg/s of carbon emitted as reported in chapter
6.1.3. These emissions are safely stored underground. As this process is related to carbon capturing,
minimal CO2 emissions are released. This value is well under the legal limits of release in legislated
in Australia (Parliament, 2013).
6. MASS & ENERGY BALANCES and Process Integration
6.1 Mass Balance PBD
6.1.1 Flue Gas
Lead author: Bryn Barker
Contributing author(s): Abdelrahman Abouelela
Table 6.1.1.1: Shows the composition of the CSG being combusted in the power plant
Component
Vol %
𝑥𝑐𝑜𝑚𝑝𝑜𝑛𝑒𝑛𝑡
Methane (CH4)
95.0
0.95
Ethane (C2H6)
3.5
0.035
Propane (C3H6)
1.0
0.01
Carbon dioxide (CO2)
0.3
0.003
Nitrogen (N2)
0.19964
0.0019964
Hydrogen Sulphide (H2S)
0.00036
0.0000036
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Assumptions:
•
•
•
•
Efficiency of the gas turbine is assumed to be 35%. (Brighthub Engineering, 2022)
15% excess air is fed to the turbine
N2 and CO2 are assumed to be inert (Mahmud T, 2023)
All combustion is complete combustion.
Stoichiometric calculations:
𝐶𝐻4 + 2𝑂2 → 𝐶𝑂2 + 2𝐻2 𝑂 (6.1.1.1)
𝐶2 𝐻6 + 3.5𝑂2 → 2𝐶𝑂2 + 3𝐻2 𝑂 (6.1.1.2)
𝐶3 𝐻8 + 5𝑂2 → 3𝐶𝑂2 + 4𝐻2 𝑂 (6.1.1.3)
𝐻2 𝑆 + 1.5𝑂2 → 𝐻2 𝑂 + 𝑆𝑂2 (6.1.1.4)
Table 6.1.1.2.HHV and density values (engineering ToolBox, 2003)
Component
HHV (kJ/m3)
Density (kg/m3)
Methane
37769
0.67
Ethane
66433
1.36
Propane
95830
1.88
Carbon dioxide
-
1.98
Nitrogen
-
1.21
Hydrogen sulphide
-
1.54
To determine the density of the coal seam gas the individual component densities shown in table
6.1.1.2 were multiplied by their respective molar fractions as shown in table 6.1.1.1.
The fuel’s HHV was calculated using the equation below:
𝐹𝑢𝑒𝑙 𝐻𝐻𝑉 = (𝑥𝐶𝐻4 × 𝐻𝐻𝑉𝐶𝐻4 ) + (𝑥𝐶2 𝐻6 × 𝐻𝐻𝑉𝐶2 𝐻6 ) + (𝑥𝐶3 𝐻8 × 𝐻𝐻𝑉𝐶3 𝐻8 ) (6.1.5)
𝐹𝑢𝑒𝑙 𝜌 = (𝑥𝐶𝐻4 × 𝜌𝐶𝐻4 ) + (𝑥𝐶2𝐻6 × 𝜌𝐶2𝐻6 ) + (𝑥𝐶3 𝐻8 × 𝜌𝐶3 𝐻8 ) + (𝑥𝐶𝑂2 × 𝜌𝐶𝑂2 ) + (𝑥𝑁2 × 𝜌𝑁2 ) + (𝑥𝐻2𝑆 × 𝜌𝐻2𝑆 ) (6.1.6)
Using equation 6.1.1 to calculate the fuel’s HHV:
𝐹𝑢𝑒𝑙 𝐻𝐻𝑉 = (0.95 × 37.669) + (0.035 × 66.433) + (0.01 × 95.830) = 39.07 𝑀𝐽/𝑚3
Equation 6.1.2 is used similarly to calculate fuel’s density. Fuel ρ= 0.717 kg/m3
The fuel HHV is the calculated in MJ/kg using the equation shown below:
𝐻𝐻𝑉(𝑀𝐽/𝑘𝑔) =
𝐻𝐻𝑉 (𝑀𝐽/𝑚3 )
𝜌 (𝑘𝑔/𝑚3 )
39.07
= 0.717 = 54.49 𝑀𝐽/𝑘𝑔 (6.1.7)
The proposed power plant wants to generate 100 MJ/s using a gas turbine. Gas turbines have an
efficiency of 0.35 (Pashchenko et al., 2022). Therefore, the thermal power is calculated using the
equation below:
𝑇ℎ𝑒𝑟𝑚𝑎𝑙 𝑝𝑜𝑤𝑒𝑟 =
𝐸𝑙𝑒𝑐𝑡𝑟𝑖𝑐𝑎𝑙 𝑝𝑜𝑤𝑒𝑟
η
100
= 0.35 = 285.71 𝑀𝐽/𝑠 (6.1.8)
The mass flow of fuel needed to be combusted to give off the required thermal power is calculated
using the equation below:
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
𝑅𝑒𝑞𝑢𝑖𝑟𝑒𝑑 𝐹𝑢𝑒𝑙 𝑖𝑛𝑙𝑒𝑡 𝑚𝑎𝑠𝑠 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 =
𝑇ℎ𝑒𝑟𝑚𝑎𝑙 𝑝𝑜𝑤𝑒𝑟
𝐻𝐻𝑉
=
285.71
54.49
= 5.243 𝑘𝑔/𝑠 (6.1.9)
To determine the mass of each component in the flue, a 100 moles basis was used to determine the
mass fractions of each component. To determine the mass of each component the formula below
was used:
𝑚𝑎𝑠𝑠 = 𝑚𝑜𝑙 × 𝑀𝑅 (6.1.10)
95 × 16 = 1520 𝑔
After the masses of each component were found, the total mass was determined for the 100 moles
basis, then the mass of each component was divided by this total mass to determine their individual
mass fractions, the results of this are shown in table 6.1.1.3:
Table 6.1.1.3. Calculated mass of each component in the fuel
These mass fractions were then multiplied by the determined flue gas mass of 5.2435 kg/s to
determine the mass flowrates of the flue gas components, these masses were then used find the
molar flowrates of the flue gas components based on the actual mass flowrate.
Table 6.1.1.4. mass and molar flowrates of fuel components
fuel
mass frac
Mass
flow
rate Molar flow rate (mol/s)
(kg/s)
CH4
0.900579485
4.722161217
295.1351
C2H6
0.062211083
0.326201926
10.8734
C3H8
0.026069406
0.136694141
3.106685
CO2
0.007820822
0.041008242
0.932006
N2
0.003311952
0.017366121
0.620219
H2S
7.25203E-06
3.80258E-05
0.001118
using the individual flue gas molar flowrates in table 6.1.1.4 and substituting them into the
stoichiometric combustion equations the flue gas compositions were calculated, a sample calculation
is shown below for this:
295𝐶𝐻4 + 590𝑂2 → 295𝐶𝑂2 + 590𝐻2 𝑂
after the stoichiometric calculations were carried out the excess air was calculated, the sum of the
stoichiometric oxygen was first calculated before this was then multiplied up to find the excess air.
these calculations are shown below.
1.15 × ∑ 𝑠𝑡𝑜𝑖𝑐ℎ𝑖𝑜𝑚𝑒𝑡𝑟𝑖𝑐 𝑂2 = 𝑠𝑢𝑝𝑝𝑙𝑖𝑒𝑠 𝑂2 (6.1.1)
𝑠𝑢𝑝𝑝𝑙𝑖𝑒𝑑 𝑂2 ÷ 0.21 = 𝑡𝑜𝑡𝑎𝑙 𝑎𝑖𝑟 𝑓𝑒𝑒𝑑 (6.1.12)
𝑡𝑜𝑡 𝑎𝑖𝑟 𝑓𝑒𝑒𝑑 × 0.79 = 𝑁2 𝑓𝑒𝑒𝑑 (6.1.13)
𝑠𝑢𝑝𝑝𝑙𝑖𝑒𝑑 𝑂2 − ∑ 𝑠𝑡𝑜𝑖𝑐ℎ𝑖𝑜𝑚𝑒𝑡𝑟𝑖𝑐 𝑂2 = 𝑂2 𝑖𝑛 𝑓𝑙𝑢𝑒 𝑔𝑎𝑠 (6.1.14)
𝑁2 𝑓𝑒𝑒𝑑 + 𝑁2 𝑖𝑛 𝑓𝑢𝑒𝑙 = 𝑁2 𝑓𝑙𝑢𝑒 𝑔𝑎𝑠 (6.1.15)
These calculations alongside the calculations for flue gas composition were used to calculate the
values shown in table 6.1.1.5:
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Table 6.1.1.5. flue gas compositions
flue gas
CO2
H2O
O2
N2
SO2
total
Molar flow
rate (mol/s)
327.13
635.32
96.58
2786.09
0.0011
3845.1206
Composition
(%)
8.51
16.52
2.51
72.46
0.29ppm
100
Mass flow rate
(kg/s)
14.394
11.436
3.091
78.011
0.000071616
106.931
Composition
(%)
13.46
10.69
67 ppm
2.89
72.95
100
6.1.2 Direct Contact Cooler (DCC-101)
All streams highlighted with yellow are streams which were unknown and calculated.
Lead author: Sari Miari
3
2
2B
2B
D-101
2
1
Table Stream table for unit DCC-101
Stream No.
1
Component (kg/s)
CO2
14.394
H2O
11.436
O2
3.091
N2
78.011
Total flow (kg/s)
106.931
Temperature (°C)
140
Pressure (bar)
1.2
Phase
V
2B
2
3
0
574.203
0
0
574.203
23
3
L
0.089
574.901
0
0
574.990
40
1.1
L
14.305
10.738
3.091
78.011
106.144
40
2.9
V
Assumptions:
•
•
•
SO2 mass flow rate is very small. Thus it is neglected and considered as 0 for the calculations.
Pressure drop is 0.1 bar only and is assumed to not affect calculations.
Only CO2 is absorbed by the cooling water, N2 and O2 absorption are neglected due to being
too small.
Firstly, the required mass flow rate of inlet cooling water was calculated using an integrated mass
& energy balance equation which is explained in section 6.2.2. The cooling water inlet flow rate
was calculated to be 31,900.182 mol/s = 574.203 kg/s.
Secondly, the amount of water vapour in the flue that condenses due to cooling of the flue gas was
calculated using the water vapour partial pressure and saturated pressure. This is shown below:
𝑝𝐻2 𝑂 = 𝑃𝑇𝑜𝑡𝑎𝑙 × 𝑥𝐻2 𝑂 = 1.2 × 0.1652 = 0.1983 𝑏𝑎𝑟 (6.1.2.1)
The value of 𝑥𝐻2 𝑂 was obtained from table 6.1.1.5 from section 6.1.1.
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
The Antoine’s equation was then used to calculate the saturated pressure of water vapour at 40 ºC
(313 K), which is the temperature of the flue gas outlet stream. The parameters for Antoine’s
equation were at 313 K were obtained from NIST, 2023.
𝐵
1733.926
log10 𝑃𝑠𝑎𝑡 = 𝐴 − (𝑇+𝐶 ) = 5.20389 − (313+(−39.485)) = −1.1355 (6.1.2.2)
𝑃𝑠𝑎𝑡 = 101.1355 = 0.0732 (6.1.2.3)
The mole fraction of the flue gas’ water vapour which will condense was calculated using the
equation shown below:
𝑃𝑠𝑎𝑡
𝑥𝐻2 𝑂𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑖𝑛𝑔 = 𝑃
𝑇𝑜𝑡𝑎𝑙
=
0.0732
1.2
= 0.061 (6.1.2.4)
The amount of water vapour in the flue gas which condenses due to cooling was finally calculated
using the equation shown below:
𝑛̇ 𝐻2 𝑂𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑖𝑛𝑔 = 𝑛̇ 𝐻2 𝑂 × 𝑥𝐻2 𝑂𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑖𝑛𝑔 = 635.318 × 0.061 = 38.751 𝑚𝑜𝑙/𝑠 (6.1.2.5)
Finally, the amount of CO2 which gets absorbed from the flue gas onto the cooling water was
calculated using Henry’s Law as shown in the equation below:
𝑝𝐶𝑂2 = 𝑥𝐶𝑂2 × 𝑃𝑇𝑜𝑡𝑎𝑙 = 0.0851 × 1.2 = 0.102 𝑏𝑎𝑟 (6.1.2.6)
𝐶 = 𝑘𝐻 × 𝑝𝐶𝑂2 = 0.03445 × 0.102 = 0.003517 𝑚𝑜𝑙/𝐿 (6.1.2.7)
The kH value for CO2 at 23 ºC was obtained from NIST, 2023. This is because Henry's law is typically
applied to the temperature of the liquid phase, which, in this case, would be the cooling water inlet
stream which is at 23 °C. The obtained value of 0.003517 mol/L means that every litre of water in
the cooling stream absorbs 0.003517 moles of CO2, which is equivalent to 0.155 grams of CO2.
Based on the calculation in 6.2.2, the molar flow rate of the cooling water is 31,900.182 mol/s =
574.203 kg/s. the specific volume of water at 23 °C is 1.00246 L/kg (ref), this is used to calculate
the volumetric flow rate of the cooling water stream:
𝑉̇ 𝐻2 𝑂𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑖𝑛𝑙𝑒𝑡 = 𝑚̇𝐻2 𝑂𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑖𝑛𝑙𝑒𝑡 × 𝑣𝑎𝑡 23 °C = 547.203 × 1.00246 = 575.616 𝐿/𝑠 (6.1.2.8)
The volumetric flow rate of water is then multiplied by absorption concentration as shown below:
𝑛̇ 𝐶𝑂2,𝐴𝑏𝑠𝑜𝑟𝑏𝑒𝑑 𝑏𝑦 𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟 = 𝐶 × 𝑉̇ 𝐻2 𝑂𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑖𝑛𝑙𝑒𝑡 = 0.0003517 × 575.616 = 2.025 𝑚𝑜𝑙/𝑠 (6.1.2.8)
2.025 mol/s of CO2 is equivalent to 0.089 kg/s.
6.1.3 Packed Bed Absorption Column (A-101)
Lead author: Bryn Barker
Contributing author(s): Faisal Kamal
36
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Stream table for unit A-101
Stream No.
3A
Component
flow
(kg/s)
14.39
CO2
H2O
10.74
3.09
O2
N2
78.01
MEA
0
106.23
Total flow (kg/s)
Temperature (°C)
40
1.2
Pressure (bar)
Phase
V
5
4
12
25.91
167.62
0
0
71.84
265.37
?
1.1
L
1.44
10.74
3.09
78.01
Trace
93.28
40
1.1
V
12.95
167.62
0
0
71.84
252.42
40
1.2
L
*Yellow values are those determined by mass balance
Assumptions:
Only carbon dioxide is absorbed by the lean MEA stream, stream 12. (Mahmud T, 2023)
90% of the CO2 from 3A is absorbed by 12 as per the briefing document
For the mass balance the vaporisation of MEA in stream 4 is negligible
From the assumption that only CO2 is absorbed by the lean MEA stream (stream 12) all of the N2,
H2O and O2 present in the flue gas from the DCC in stream 3A will leave the absorption column in
stream 4. Since 90% of the CO2 is absorbed by stream 12 then the CO2 present in stream 4 was
given by equation 6.1.3.1 shown below,
14.39𝑚̇ × 0.1 = 1.44𝑚̇𝐶𝑂2 (6.1.3.1)
From this calculation the composition of stream 4 was fully determined. From the mass of CO2
present in the flue gas stream the mass of CO2 absorbed by the MEA stream inside the absorption
unit was determined, this calculation is shown in equation 6.1.3.2 below.
14.39𝑚̇ × 0.9 = 12.95𝑚̇𝐶𝑂2 (6.1.3.2)
eq 6.1.3.2
the loading values of the lean and rich MEA streams were set as 0.25 moles CO2/mole MEA and 0.5
moles CO2/mole MEA respectively. From these loading values the MEA required for the absorption
of the calculated CO2 was calculated . the equations used in these calculations were as follows:
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
0.5 𝑚𝑜𝑙 𝐶𝑂2
𝑚𝑜𝑙 𝑀𝐸𝐴
−
0.25 𝑚𝑜𝑙 𝐶𝑂2
𝑚𝑜𝑙 𝑀𝐸𝐴
=
0.25 𝑚𝑜𝑙 𝐶𝑂2
𝑜𝑓
𝑚𝑜𝑙 𝑀𝐸𝐴
𝑎𝑑𝑑𝑖𝑡𝑖𝑜𝑛𝑎𝑙 𝑙𝑜𝑎𝑑𝑖𝑛𝑔 (6.1.3.3)
this additional loading is equal to the molar amount of CO2 that is required to be absorbed in the
absorber, this molar amount is calculated from the mass flowrate determined in equation 6.1.3.2.
consequently the number of moles of CO2 per second can be multiplied by the additional loading
value calculated in equation 6.1.3.3 to determine the number of moles present in both streams 12
and 5. This calculation is shown in equation 6.1.3.4 below
294.32𝑛̇ 𝐶𝑂2 × 4 = 1177.28𝑛̇ 𝑀𝐸𝐴 (6.1.3.4)
eq 6.1.3.4
From these calculations the moles of CO2 present in streams 12 and 5 can be calculated from the
molar flowrate of MEA and the pre-determined loading values, the calculations are shown in
equations 6.1.3.5 and 6.1.3.6 below.
0.25 × 1177.28 = 294.32𝑛̇ 𝐶𝑂2 𝑖𝑛 𝑠𝑡𝑟𝑒𝑎𝑚 4 (6.1.3.5)
0.5 × 1177.28 = 588.64𝑛̇ 𝐶𝑂2 𝑖𝑛 𝑠𝑡𝑟𝑒𝑎𝑚 2 (6.1.3.6)
The process selection determined that the optimum concentration of MEA in an MEA water solution
was 30wt%. From this the molar flowrate of MEA calculated in equation 6.1.3.4 was converted into
a mass flowrate and this was used to calculate the mass flowrate of water required in streams 5 and
12. This was determined to be 167.62 KG/s of water
6.1.4 Water wash (w-101)
Lead author: Bryn Barker
Stream table for unit w-101
Stream No.
Component
flow
(kg/s)
CO2
H2O
O2
N2
MEA
Total flow (kg/s)
Temperature (°C)
4
4A
wwin
wwout
1.44
10.74
3.09
78.01
0.0093
93.37
40
1.44
10.74
3.09
78.01
0.00016
93.28
40
0
1.35
0
0
0
1.35
23
0
1.35
0
0
0.0091
1.36
23
38
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Pressure (bar)
Phase
1.1
V
1.1
V
1.1
L
1.1
L
*Note yellow values are those calculated by mass balance
Assumptions:
The water wash is capable of reducing the MEA emissions in the flue gas to 1.7 ppm (Liu et al.,
2021)
The MEA Concentration in the water wash water after absorption is 0.2 mol% (Liu et al., 2021)
only MEA is absorbed by the water wash water
Stream 4 is the cleaned flue gas from the absorber as calculated in mass balance 6.1.3. The vapour
pressure of the liquid MEA was first determined by the following equation using 313K (40°C) for T
as this was the temperature of the absorption column tops. (Nguyen et al., 2010)
−1.04×104
)−
𝑇
0
𝑃𝑀𝐸𝐴
= 𝐸𝑋𝑃 [92.6 + (
(9.47𝐿𝑛𝑇) + (1.9 × 10−18 𝑇 6 )] (6.1.4.1)
this gave an MEA vapour pressure of 0.1424KPa. This calculated vapour pressure was substituted
into equation 6.1.4.2 . (Nguyen et al., 2010)
𝛾𝑎𝑚𝑖𝑛𝑒 =
𝑃𝑎𝑚𝑖𝑛𝑒
0
𝑥𝑎𝑚𝑖𝑛𝑒 ×𝑃𝑎𝑚𝑖𝑛𝑒
(6.1.4.2)
for this equation 𝛾𝑎𝑚𝑖𝑛𝑒 , the activity coefficient of the MEA, was taken to be 1, and 𝑥𝑎𝑚𝑖𝑛𝑒 was the
mole fraction of MEA in stream 4. Rearranging equation x a gave a partial pressure of the MEA,
𝑃𝑎𝑚𝑖𝑛𝑒 , in stream 4 of 0.0048KPa. (Nguyen et al., 2010) Because the partial pressure of the MEA
vapour was determined Dalton’s law was then used to work out the number of moles of MEA present
in stream 4, Dalton’s law is shown in equation 6.1.4.3 below
𝑃𝑡𝑜𝑡𝑎𝑙
𝑛𝑡𝑜𝑡𝑎𝑙
𝑃
= 𝑛𝑎𝑚𝑖𝑛𝑒 (6.1.4.3)
𝑎𝑚𝑖𝑛𝑒
for the Dalton’s law equation 𝑃𝑡𝑜𝑡𝑎𝑙 was taken to be 110Kpa, this was determined from the pressure
drop in the absorber, from this goal seek was used to determine the moles of amine present in
stream 4. The equation used to calculate 𝑛𝑎𝑚𝑖𝑛𝑒 is shown in equation 6.1.4.4.
110
3511.86+𝑛𝑎𝑚𝑖𝑛𝑒
0.0048
=𝑛
𝑎𝑚𝑖𝑛𝑒
(6.1.4.4)
the number of moles of MEA in stream 4 was determined to be 0.15mol/s. This was calculated to be
99.6ppm. From this calculation as stated in the assumptions the concentration had to be reduced to
1.7ppm. From these calculations it was determined that 9.13g/s of MEA would have to be removed
from the flue gas by the water wash. From the assumption that the water wash should contain 0.2
39
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
mol% MEA it was calculated that the water wash would require a water flowrate of 1.35kg/s to
remove the MEA.
6.1.5 S-101
Lead author: Faisal Kamal
Contributing author(s): Faisal Kamal
Due to lack of initial information and the “delaying” of the calculation of the column stages, it is
assumed that the stream 6 is in equilibrium with the stream 5B, for the same reasons, a “stripping
split” was assumed between the column itself and the reboiler, The Ratio is 70% of CO2 Stripping is
done in the column while 30% occurs in the reboiler. It is also assumed that both the top and bottom
product streams are in equilibrium with the Reflux and Boil up streams respectively. The pressure in
the system is also assumed to be constant.
6
CON-101
5B
8
7
S-101
10
REB-101
9
11
Table 6.1.5.1: Stream
Stream No.
CO2 flow (Kg/s)
MEA flow (Kg/s)
H2O Flow (Kg/s)
Total flow (Kg/s)
Temperature (°C)
Pressure (bar)
Phase
5B
25.91
71.83
167.71
265.47
102.78
1.2
L
6
?
?
?
?
102.78
1.8
V
table for unit S-101
7
8
9
?
12.95
?
?
?
?
?
?
?
?
?
?
40
40
?
1.8
1.8
2
L
V
L
10
?
?
?
?
120
2
V
11
12.95
?
?
?
120
2
L
The aim of this mass balance is to determine the flow rates of each stream, the ultimate aim is to
determine the purity of the of the top product, stream 8. The temperature of stream 6 is known from
the assumption of equilibrium with the Rich MEA stream inlet, therefore thermal equilibrium is an
extension of that assumption. Furthermore, the pressure and temperature of all units are set in
accordance with literature as set in chapter 3, these pressure are chosen with the best performance
of the solvent used and loading values in mind. The temperatures of streams 7 and 8 are set to the
temperature of the condenser, while streams 10 and 11 are set to the temperature of the reboiler.
The amount of CO2 leaving at stream 8 is a set value as it is the difference of the lean and rich
loading values of 0.25 and 0.5 mol/ mol MEA.
40
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6.1.5.1 Condenser mass balance:
6.1.5.1.1 Stream 6:
The first step of the mass balance is to calculate the concentrations of Stream 6. Since equilibrium
with stream 5B was assumed. The calculation would follow simple bubble point, dew point, flash
calculations (Coulson and Richardson, 2005). Therefore, to calculate the concentrations of stream
6 we employ the following equation:
𝑦𝑖 = 𝐾𝑖 𝑥𝑖 (6.1.5.1)
In this case, yi would be the concentration of a component in stream 6, while the xi value would
signify the molar concentration of a stream 5B component. The Ki value would be the equilibrium
constant of said component at the specific temperature that both streams share.
This constant is acquired by dividing the Antione’s equation of a component by the unit’s pressure
and is done at the temperature of the of the stream. This equilibrium constant is key to determining
the concentrations of the components in Stream 6. The equation for Ki is:
𝐾𝑖 =
10
𝐵
𝐴−
𝑇+𝐶
𝑃
(6.1.5.2)
For Example, in determining vapour concentration of H2O, we must first define xH2O which would be:
𝑥𝐻2𝑂 =
𝑚𝑜𝑙
)
𝑠
𝑚𝑜𝑙
𝑟𝑎𝑡𝑒 (11078.93
)
𝑠
𝑀𝑜𝑙𝑎𝑟 𝑓𝑙𝑜𝑤 𝑟𝑎𝑡𝑒 𝑜𝑓 𝐻2 𝑂 (9312.41
𝑇𝑜𝑡𝑎𝑙 𝑀𝑜𝑙𝑎𝑟 𝑓𝑙𝑜𝑤
= 0.8405 (6.1.5.3)
The equation was repeated to attain other concentrations where 𝑥𝐶𝑂2 =0.053 and 𝑥𝑀𝐸𝐴 = 0.106.
Finally, to determine the equilibrium constant KH2O, Antoine’s coefficients for H2O must be used at
the appropriate temperature, these can be seen in table 1 below:
Table 6.1.5.2: Antoine's coefficients of MEA and H2O, where the pressure is in (Bar) and the
temperature in (Kelvin)
MEA
H2O
TEMPERATURE
(K)
338.6-444.1
344 - 373
RANGE
A
B
C
4.29252
5.08354
1408.873
1663.125
-116.093
-45.622
With these parameters, KH2O value can be calculated as follows:
1663.125
𝐾𝐻2𝑂
105.08354−375.78−45.622
=
= 0.617
1.8
The equations 6.1.5.2 and 6.1.5.3 are repeated again for MEA with the stream 5B’s liquid
concentration and the MEA Antoine’s coefficients respectively, this yields 𝐾𝑀𝐸𝐴 = 0.0409. Equation
6.1.5.3 is repeated for CO2 This allows the calculation of the vapour concentrations of these two
components using Equation (6.1.5.1). For example, in the case of H2O, the vapour concentration
would be as follows:
41
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
𝑦𝐻20 = 0.617 ∗ 0.8405 = 0.519
The repetition of the previous three equations to determine the value the vapour concentration of
MEA in stream 6 yields the value of yMEA = 0.0021. This value is considerably smaller than the water
value due to MEA having the lowest volatility of all the components. Having obtained both these
values, we can calculate the concentration CO2 in stream 6 to be yCO2 = 0.478. This is done as the
summation of all concentrations in a compound is equal to 1:
𝛴𝑥𝑖 𝑜𝑟 𝛴𝑦𝑖 = 1
Equation (6.1.5.4)
Notably, KCO2 calculations has not been conducted in stream 6. This is because the mass balance
calculation does not require the specific calculation of the equilibrium constant of CO 2 at this stage.
The second rationale is that CO2 has a critical temperature of 31°C. Consequently, at temperatures
higher than that, Antoine's equations are unavailable and are less accurate. This will be covered in
more detail in the stream 7 calculations.
6.1.5.1.2 Stream 7:
Stream 3 is the reflux stream of the condenser; this stream can be calculated using the information
acquired from the calculations of stream 6 above. The concentrations of stream 7 can be obtained
by:
𝑥𝑖 =
𝑧𝑖
1 − 𝑉(𝐾𝑖 − 1)
Equation (6.1.5.5)
In this equation the 𝑧𝑖 is the inlet concentration, in this case this is the vapour concentrations 𝑦𝑖 from
stream 6. V is the vapour fraction in the condenser, i.e., how much of stream 6 vaporises. This value
was assumed at 0.5 to achieve initial results. However, through goal seeking a value of the vapour
fraction was attained that would satisfy equation (6.1.5.4), where the vapour fraction value attines
sum of 𝑥𝑖 to equal 1.
Notably, 𝐾𝑖 must be calculated again for stream as its temperature is lower (40°C), this would require
new Antoine’s coefficient, these, and all other coefficients are taken from national institute of
standers and technologies (NIST, 2023). These can be seen in table 2 below:
Table 6.1.5.3: Antione's equations at the range of 40°C, temperature and pressure of CO2 and MEA
are in Celsius and mmHg. H2O temperature and pressure are at kelvin and Bar
MEA
H2O
CO2
TEMPERATURE RANGE
(K)
283-637
303-334
216-304
A
B
8.36214
5.20389
7.8101
2117.92
1733.926
987.44
C
215.389
-39.485
290.9
In the case of CO2 and MEA, to perform the calculations correctly, the temperature in to be used
Antoine’s equation was converted into Celsius, the pressure value from Antoine’s equations was
42
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
turned into bar by dividing it by 750.1 before being dived by the unit’s pressure to attain the
equilibrium constant. For example, in the case of MEA the calculation would be:
2117.92
108.36214−40+215.389
(
)
750.1
𝑘𝑀𝐸𝐴 =
= 0. 000868621
1.8
This method is repeated again to calculate the KCO2 Values. It’s important to note that the temperature
range for CO2 in the above table is the highest range possible for CO2 as mentioned before 31°C is
the critical temperature for CO2. This was done in accordance with advice form consultants, the
optimum way of attaining the KCO2 at these conditions would be software such as Aspen Plus. As
access to Aspen is not granted at this stage of the calculations, and as the temperature difference
between the condenser and the temperature limit of Antoine’s coefficients is relatively small at 9 °C.
A decision was made to use the coefficients at the critical temperature of CO2.
As Antione’s parameters for H2O were in the units Kelvin and Bar, no conversion was needed before
or during the calculations.
This information is all that is required to calculate the concentrations in stream 7. For instance, as
per equation (6.1.5.5), in the case of MEA:
𝑥𝑀𝐸𝐴 =
0.0021
= 0.004244
1 − 0.487924927(0. 000868621 − 1)
This equation was repeated two additional times with the values of H2O and CO2 to calculate their
liquid molar concentrations in stream 7. Where it was found that 𝑥𝐶𝑂2 = 0.01935 and 𝑥𝐻2𝑂 = 0.9764.
It should be noted that the vapour fraction used in the equation above and in the calculations for the
𝑥𝑖 values is the vapour fraction value that was obtained using goal seeking to ensure that the sum
of 𝑥𝑖 is equal to 1. The results of the 𝑥𝑖 values summed can be used to confirm the success and
accuracy of the vapour fraction value used:
0.019355276 + 0.0042448 + 0.9764 = 1
6.1.5.1.3 Stream 8:
In stream 8 the flow rate of CO2 is already specified; this means that calculating all the concentrations
in the stream will provide the rest of the flow rates in the stream. In order to calculate the
concentration, equation 6.5.1.1 is repeated here, where 𝑥𝑖 is equal to the previously calculated liquid
concentrations in stream 7. The Ki are also the same as the temperature and pressure are the same
in both streams. Therefore, for example, in the case of stream 8’s MEA:
𝑦𝑀𝐸𝐴 = 0.000868621 ∗ 0.0042448 = 0.0008481
As mentioned before, small concentrations of MEA are expected in all vapour phases due to MEA’s
low volatility relative to other components in the system. This can be seen with the rest of the
calculation of the concentrations as 𝑦𝐻2𝑂 = 0.0397 and 𝑦𝐶𝑂2 = 0.95944. The value of 𝑦𝐶𝑂2 signifies
43
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
that Stream 8 has a CO2 purity of 95.94%, this is deemed to be appropriate and is consistent with
literature (Goto et al., 2013)
6.1.5.1.4 Flowrates around condenser:
Having made all preliminary calculations, the flow rates in the above streams can know be
calculated. Firstly, in stream 8, the flow rate of CO2 is known, therefore the flowrate of the two other
components in the stream can be calculated. However, to start, the overall flow rate in stream 8 must
be calculated first using:
𝑂𝑣𝑒𝑟𝑎𝑙𝑙 𝑀𝑜𝑙𝑎𝑟 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 𝑛̇ =
𝑛̇ 𝑖
𝑦𝑖
Equation (6.1.5.6)
In this case, 𝑛̇ 𝑖 is the CO2 molar flowrate which is already to be 294.45 mol/s (Table 1). Similarly, 𝑦𝑖
is the molar fraction of CO2 in the composition previously calculated at 0.95944. Therefore, applying
equation (6.1.5.6) on these values yields and overall molar flow of 306.86 mol/s. Having calculated
the overall flow rate in the stream, it is possible now to define the flowrates of the two other
components:
𝑛̇ 𝑖 = 𝑛̇ ∗ 𝑦𝑖
Equation (6.1.5.7)
Notably, in both pervious equations, 𝑦𝑖 or 𝑥𝑖 can be used depending on the state of the stream in
question.
Using the equation (6.1.5.7), the molar flow rates of both H2O and MEA were calculated at 12.18
mol/s and 0.26 mol/s respectively. In converting these flowrate values, CO2 = 12.95 Kg/s, H2O =
0.21 Kg/s, and MEA= 0.0158 Kg/s. The calculation of these values were the last remaining unknowns
in stream 8. Having these values allows for the calculation of the flowrates of the other two streams
before it.
Consequently, the values from stream 8 alongside the vapour fraction used in equation (6.5.1.3) can
be used to calculate the amount of CO2 in the inlet to the condenser Stream 6. The vapour fraction
is an indication of split into vapour or liquid in the condenser. For example, in the case of CO2, V =
0.487924927 is equal to 12.95 Kg/s, this is because 48.79% of the CO2 vaporised, or in this case
did not condense and stayed in the vapour phase, and from that value, the amount of overall CO2
coming in from the inlet cane be calculated. Therefore, equation (6.1.5.6) can be used here to
determine the flowrate of CO2 in stream 6:
𝑚𝑜𝑙
294.42 ( 𝑠 )
𝑚𝑜𝑙
𝑆𝑡𝑟𝑒𝑎𝑚 2 𝐶𝑂2 𝑀𝑜𝑙𝑎𝑟 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 =
= 603.41 (
)
0.487924927
𝑠
With this value obtained, alongside the 𝑦𝑖 values from Stream 6 Calculations, the overall molar
flowrate of Stream 6 can be calculated using the same equation. Thus, the overall molar flowrate in
stream 6 is 1261.15 mol/s. Using equation (6.1.5.7), the rest of the molar flowrates of the individual
44
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
components are calculated at; 654.99 mol/s for H2O, and 2.743 mol/s for MEA. In converting these
values, in stream 6: MEA = 0.167 Kg/s, H2O= 11.79 Kg/s, and CO2= 26.55 Kg/s.
Finally, stream 7 cane be calculated using simple mass balance, where it would be equal to the
difference between stream 6 and stream 8:
Equation (6.1.5.8)
𝑆𝑡𝑟𝑒𝑎𝑚 7 = 𝑆𝑡𝑟𝑒𝑎𝑚 6 − 𝑆𝑡𝑟𝑒𝑎𝑚 8
And so, the streams around the condenser are equal to:
Table 6.1.5.4: Table showing flowrates of streams 6 to 8.
Stream No.
6
7
8
CO2 flow (Kg/s)
26.55623271
13.59878481
12.9574479
MEA flow (Kg/s)
0.167329041
11.57704016
0.01587577
H2O Flow (Kg/s)
11.79646581
0.151453271
0.219425647
Total flow (Kg/s)
38.52002756
25.32727825
13.19274932
Temperature (°C)
102.78
40
40
Pressure (bar)
1.8
1.8
1.8
Phase
V
L
V
6.1.5.2 Mass Balance around Reboiler and column:
6.1.5.2.1 Stream 11
Stream 11 is the outlet stream from the reboiler, this stream is known as the Lean solvent stream.
As this is one of the only two outlets around the entire system, the other being stream 8. The
compositions and flowrates of the stream can easily be determined using simple in = out mass
balance, and therefore:
𝑆𝑡𝑟𝑒𝑎𝑚 11 = 𝑆𝑡𝑟𝑒𝑎𝑚 5𝐵 − 𝑆𝑡𝑟𝑒𝑎𝑚 8
Equation (6.1.5.9)
Subtracting the flow rates of the pure CO2 stream, stream 8, and the Rich solvent inlet stream, stream
5B, yields the following values for each component:
𝑆𝑡𝑟𝑒𝑎𝑚 11 𝐶𝑂2 = 25.91 − 12.95 = 12.96
𝐾𝑔
𝑠
𝑆𝑡𝑟𝑒𝑎𝑚 11 𝑀𝐸𝐴 = 71.83 − 0.015 = 71.78
𝐾𝑔
𝑠
𝑆𝑡𝑟𝑒𝑎𝑚 11 𝐻2 𝑂 = 167.71 − 0.21 = 167.50
𝐾𝑔
𝑠
The total flowrate in stream 11 is 252.24 Kg/s. Notably, as the values of stream 8 are the difference
between the Rich and lean solvents, the lean solvent must be fed these differences as “make up
Water/MEA”. This compensation of lost material has been further discussed in Chapter 3.
45
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6.1.5.2.2 Stream 10
Stream 10 contains the vapour produced in the reboiler, hence it is known as the boil up stream.
Based on the aforementioned assumption that 30% of CO2 stripping occurs in the reboiler, the CO2
in stream 10 is 30% of that in stream 6. Thus, stream 6 contains 7.96 Kg/s or 181 mol/s of CO2.
Furthermore, since stream 10 is in equilibrium with stream 11, equations (6.1.5.1-6.5.1.4) can be
used in this case to calculate the vapour concentrations in stream 10 using the 𝑥𝑖 concentrations of
stream 11 and the 𝐾𝑖 values at the temperature of the reboiler, 120 °C.
Notably, Because MEA in both streams 10 and 6 is in the appropriate temperature range, the
Antione's coefficients used for MEA are the same as those in Table 1. But since H2O is outside of
Table 1's temperature range, new Antione's coefficients must be applied. Table X below shows the
new H2O coefficients. Additionally, the reboiler's pressure is 2 Bar, so the 𝐾𝑀𝐸𝐴 value will still differ.
Table 6.1.5.5: Antone's coefficients of MEA at 120 °C, Pressure in Bar, Temperature in °C.
TEMPERATURE RANG (K)
A
B
C
H2O
379-573
3.55959
643.748
-198.043
From this, the 𝐾𝑖 values of two of the components can be attained. Where 𝐾𝐻2𝑂 = 0.9048 and,
𝐾𝑀𝐸𝐴 = 0.0800. 𝐾𝐶𝑂2 is not calculated for the same reasons mentioned in previously.
Having calculated the 𝐾𝑖 values, the 𝑥𝑖 values can be obtained from Stream 11 using equation
(6.1.5.3), where it can be found that, 𝑥𝑀𝐸𝐴 = 0.1093 , 𝑥𝐻2𝑂 = 0.8633 , and 𝑥𝐶𝑂2 = 0.0273 . In
acquiring both the equilibrium constant and the liquid concentration, the vapour concentrations in
stream 10 can be calculated using equation (6.1.5.1), where it can be found that 𝑦𝐻2𝑂 = 0.7811 and
𝑦𝑀𝐸𝐴 = 0.0087, and, using equation (6.1.5.4) 𝑦𝐶𝑂2 = 0.2102.
With the concentrations calculated in stream 10, the flowrates can be easily calculated as the
flowrate of CO2 is known. Therefore, using equation (6.1.5.6), the overall flowrate in stream 6 is:
𝑚𝑜𝑙
181 (
)
𝑠 = 861.776 𝑚𝑜𝑙 𝑜𝑟 20.52 𝐾𝑔
𝑆𝑡𝑟𝑒𝑎𝑚 10 𝑂𝑣𝑒𝑟𝑎𝑙𝑙 𝑀𝑜𝑙𝑎𝑟 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒 =
0.2102
𝑠
𝑠
From this, using equation (6.1.5.7), the rest of the flowrates can be calculated:
𝑆𝑡𝑟𝑒𝑎𝑚 10 𝑛𝐻20 = 0.7811 ∗ 861.776 = 673.20
𝑚𝑜𝑙
𝐾𝑔
𝑜𝑟 12.12
𝑠
𝑠
𝑆𝑡𝑟𝑒𝑎𝑚 10 𝑛𝑀𝐸𝐴 = 0.0087 ∗ 861.776 = 7.544
𝑚𝑜𝑙
𝐾𝑔
𝑜𝑟 0.46
𝑠
𝑠
6.1.5.2.3 Stream 9
The calculations in stream 9 are a simple mass balance similar to that of stream 11, where we can
take IN = OUT around the reboiler, where out is streams 11 and 10, and IN is stream 9 itself.
Therefore, to calculate stream 9:
46
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
𝑆𝑡𝑟𝑒𝑎𝑚 9 = 𝑆𝑡𝑟𝑒𝑎𝑚 11 + 𝑆𝑡𝑟𝑒𝑎𝑚 10
Equation (6.1.5.10)
And so, the flowrates for each component in stream 9 can be calculated using equation (6.1.5.7)
and are as follows:
𝑛𝑐𝑜2 = 475.44
𝑚𝑜𝑙
𝐾𝑔
𝑜𝑟 20.92
𝑠
𝑠
𝑛𝑀𝐸𝐴 = 1184.96
𝑛𝐻2𝑂 = 9973.43
𝑚𝑜𝑙
𝑠
𝑜𝑟 179.62
𝑚𝑜𝑙
𝐾𝑔
𝑜𝑟 72.28
𝑠
𝑠
𝐾𝑔
.
𝑠
6.1.5.2.4 Flow rates around the reboiler and column.
All stream flowrates have already been calculated, and therefore this section will be a calculation of
the changes in the column and a confirmation of the rest of the streams. Firstly, in the column, there
is an amount of CO2 vaporised as per the assumption of 70% of CO2 stripping happening in the
column. This amount can be known by getting 70% of the CO2 value in stream 6. Thus, in the column
itself, there is 422.38 mol/s or 18.59 Kg/s. This is not the only reaction in the column however, as it
is expected that H2O and MEA undergo changes in the column. It is simple to calculate these
changes with the knowledge of the rest of the streams in the system.
The change happening in the column can be calculated as the difference between the inlet to the
condenser and the boil up stream, streams 6 and 10 respectively. Taking out that difference resulted
from subtracting stream 10 to stream 6 gives the amount of the specific material
vaporised/condensed in the column:
Column 𝑚𝑎𝑠𝑠 𝑏𝑎𝑙𝑎𝑛𝑐𝑒 = 𝑆𝑡𝑟𝑒𝑎𝑚 6 − 𝑆𝑡𝑟𝑒𝑎𝑚 10
Equation (6.1.5.11)
If the difference is positive, it indicates the vaporisation of a component (i.e. CO2 vaporisation in the
column), if negative it indicated condensation for instance, in the case of H2O:
𝐾𝑔
𝐾𝑔
𝐾𝑔
− 12.12
= −0.33
𝑠
𝑠
𝑠
This shows that there are 0.33 Kg/s of H2O condensing in the column, doing the same calculation
11.79
for MEA, yields a value of -0.29 Kg/s. This means both MEA and water condense in the column while
only CO2 evaporates. This discrepancy between the MEA and H2O condensation and the CO2
vaporisation, is again due to the volatility of each component, where CO2 has the lowest.
Furthermore, the condensation and vaporisation occur in the column due to the heat transfer
happening between the stream 7, the reflux stream, and stream 10, the boil up stream. As the hot
vapour proceeds to the top of the column, it is cooled down by the colder reflux stream. Meanwhile
stream 7 and the inlet stream 5B are heated up by the high temperature of the boil up, stream 10,
coming out of the reboiler, this heating vaporises the CO2 in streams 5B and 7.
These values could help further corporate the values of stream 9. This can be done by subtracting
vaporisation of a component or adding the condensation of a component to stream 5B + stream 7.
Thus, for each component in stream 9:
Equation (6.1.5.11)
47
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𝑆𝑡𝑟𝑒𝑎𝑚 9 = 𝑆𝑡𝑟𝑒𝑎𝑚 5𝐵 + 𝑆𝑡𝑟𝑒𝑎𝑚 7 + 𝐶𝑜𝑛𝑑𝑒𝑛𝑠𝑎𝑡𝑖𝑜𝑛 𝑖𝑛 𝑐𝑜𝑙𝑢𝑚𝑛 (𝑜𝑟 − Vaporisation in colmn )
Running this calculation with the above values provided of vaporisation and condensation yields the
same values previously calculated in stream 9. This is done as a way of confirming both calculations.
Therefore, the final values of all flowrates around the reboiler are as follows:
Table 6.1.5.6: mass balance around reboiler.
Stream No.
9
10
11
CO2 flow (Kg/s)
20.92431771
7.966869814
12.9574479
MEA flow (Kg/s)
72.28294379
0.460208161
71.82273563
H2O Flow (Kg/s)
179.6215949
12.12446984
167.4971251
Total flow (Kg/s)
272.8288564
20.55154782
252.2773086
Temperature (°C)
?
120
120
Pressure (bar)
2
2
2
Phase
L
V
L
6.1.5.3 Mass balance summary table
Table X below shows the final and complete mass balance around the stripper system. The table
contains the flowrates, phases, temperatures, and pressures of each stream. Furthermore, the table
contains the vaporisation and condensation flowrates occurring inside the column.
Table 6.1.5.7: Final mass balance summary table
Stream No.
CO2 flow (Kg/s)
MEA flow (Kg/s)
H2O Flow (Kg/s)
Total flow (Kg/s)
Temperature (°C)
Pressure (bar)
Phase
5B
25.910
71.830
167.717
265.470
102.780
1.200
L
6
26.556
0.167
11.796
38.520
102.780
1.800
V
7
13.599
0.151
11.577
25.327
40.000
1.800
L
8
12.957
0.016
0.219
13.193
40.000
1.700
V
9
10
20.924
7.967
72.283
0.460
179.622 12.124
272.829 20.552
110.500 120.000
1.800
2.000
L
V
11
Vap in collumn Cond in column
12.957
18.589
--------------71.823
--------------1.111
167.497
--------------0.328
252.277
18.589
1.439
120.000
----------------------------1.900
1.800
1.800
L
V
L
With these mass balance finished, the reflux and boil up rations can be calculated as, calculated as
0.32 and 0.08 respectively. These values were found to be constant with literature (Gilliland, 1940).
The mass balance is based in the assumptions made, most of the assumptions made are practically
realistic (Coulson and Richardson, 2005), however, the assumption of stream 6 being in equilibrium
with stream 5B is not entirely based in reality. The actual temperatures will differ in these streams
and thus, the mass balance calculations. Notably, without this assumption and without knowledge
on the design of the column, such as column diameter, and column height, calculating the mass
balance across the systems is not feasible. Furthermore, the assumption regarding the amount of
stripping occurring in the column or reboiler are an estimation of reality and are not to be taken as a
fact. The precision of these calculations is likely to be affected should the design parameters of the
column be known (Coulson and Richardson, 2005). The number of stages in particular would lead
to more precise and closer to reality calculations. This method followed is the typical method used
when the design parameters of the column are not known, therefore, it provides a good intimal
estimation despite not being entirely accurate.
48
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6.1.6 Compressor train (COM-1 to COM-4, C-102 to C-105, D-101&D-102)
Lead author: Abdelrahman Abouelela
Contributing author(s): N/A
Mass Balance:
Assumptions:
- All Water and MEA is removed by the dryer.
The mass balance for the compressor train is very straightforward as a concentrated stream of CO2
is only compressed and cooled, with no actual reaction taking place. However, there is some water
and trace amounts of MEA in the stream, which are removed through the use of an adsorption dryer
before they enter the compressor train. The following is the mass balance for the entire process:
Table 6.1.6.1: Stream table for compressor train
Stream No.
8
13
14
12.95
12.95
12.95
0.22
0
0
0.02
0
0
13.19
12.95
12.95
Temperature
(°C)
40
40
270
Pressure (bar)
1.7
1.7
14
Phase
V
V
L
CO2 flow (kg/s)
H2O flow (kg/s)
MEA flow (kg/s)
Total flow (kg/s)
*Streams 14-21 are the same, except for temperature and pressure
49
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6.2 Energy Balance
6.2.1 Direct Contact Cooler (DCC-101)
Lead author: Sari Miari
Contributing author(s): N/A
Stream table for unit D-101
1
Stream No.
106.931
Total flow (kg/s)
140
Temperature (°C)
1.2
Pressure (bar)
V
Phase
For the energy balance:
2B
574.203
23
3
L
2
574.990
40
2.9
L
3
106.144
40
1.1
V
̂1 + 𝑛2𝐵
̂2𝐵 = ∑ 𝑛̇ 3 𝐻
̂3 + 𝑛̇ 2 𝐻
̂2
∑ 𝑛̇1 𝐻
̇ 𝐻
ṅ in (mol/s)
substance
̂ in (kJ/mol)
H
̂ 1,1 + H
̂ 1,2B
H
H2O
ṅ 1 +ṅ 2b
CO2
ṅ 1
̂2
H
O2
ṅ 1
N2
ṅ 1
̂3
H
̂4
H
ṅ 3 + ṅ 2
̂ out (kJ/mol)
H
̂ 5,3 + H
̂ 5,2
H
ṅ 3 + ṅ 2
̂ 6,3 + H
̂ 6,2
H
ṅ out (mol/s)
̂7
H
̂8
H
ṅ 3
ṅ 3
Table 6.2.2.1. cp a + bT + cT2 + dT3 kJ/mol °C
substance
H2O (g)
H2O (l)
CO2 (g)
O2 (g)
N2 (g)
a x 103
33.46
75.4
36.11
29.1
29.0
b x 105
0.688
0
4.233
1.158
0.2199
c x 108
0.7604
0
-2.887
-0.6076
0.5723
d x 1012
-3.593
0
7.464
1.311
-2.871
Heat of Vaporisation = 40.656 kJ/mol, and Heat of Condensation = -40.656
Stream 1 T= 140 °C. Stream 2B T= 23 °C. Stream 3 T=40 °C. Stream 2 T=40 °C. Tref = 25 °C
̂ 1,1 = ∫100 𝑐𝑝,𝐻 𝑂(𝑙) 𝑑𝑇 + ∆H
̂ V (100) + ∫140 𝑐𝑝,𝐻 𝑂(𝑔) 𝑑𝑇 (6.2.1.1)
H
2
2
25
100
̂ 1,1 = (75.4 × 10−3 × (100 − 25)) + 40.656 + (33.46 × 10−3 × (140 − 100)) + (1 × 0.688 × 10−5 ×
H
1
1
2
(1402 − 1002 )) + ( × 0.7604 × 10−8 × (1403 − 1003 )) + ( × −3.593 × 10−12 × (1404 − 1004 )) =
3
4
47.687 𝑘𝐽/𝑚𝑜𝑙
50
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Equations 6.2.2.2 to 6.2.2.10 are all used in the same way that equation 6.2.2.1 has been used.
̂ 1,2B = ∫23 𝑐𝑝,𝐻 𝑂(𝑙) 𝑑𝑇 = −0.151 𝑘𝐽/𝑚𝑜𝑙
H
2
25
(6.2.1.2)
̂ 5,2 = −∆H
̂ V (100) + ∫40 𝑐𝑝,𝐻 𝑂(𝑙) 𝑑𝑇 =
H
2
25
−39.525 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.7)
̂ 2 = ∫140 𝑐𝑝,𝐶𝑂 (𝑔) 𝑑𝑇 = 4.528 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.3)
H
2
25
̂ 6,3 + H
̂ 6,2 = ∫40 𝑐𝑝,𝐶𝑂 (𝑔) 𝑑𝑇 = 0.562 𝑘𝐽/𝑚𝑜𝑙
H
2
25
(6.2.1.8)
̂ 3 = ∫140 𝑐𝑝,𝑂 (𝑔) 𝑑𝑇 = 3.451 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.4)
H
2
25
̂4 =
H
140
∫25 𝑐𝑝,𝑁2 (𝑔) 𝑑𝑇
= 3.361 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.5)
40
∫25 𝑐𝑝,𝐻2 𝑂(𝑙) 𝑑𝑇
̂ 5,3 =
H
(6.2.1.6)
̂ 7 = ∫40 𝑐𝑝,𝑂 (𝑔) 𝑑𝑇 = 0.442 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.9)
H
2
25
̂ 8 = ∫40 𝑐𝑝,𝑁 (𝑔) 𝑑𝑇 = 0.436 𝑘𝐽/𝑚𝑜𝑙 (6.2.1.10)
H
2
25
= 1.131 𝑘𝐽/𝑚𝑜𝑙
Table 6.2.2.2: Shows the enthalpy
substance
H2O
CO2
O2
N2
̂ 𝐼𝑛
nIn,flue
̇ gas × H
(kJ/s)
30296.1582
1481.49665
333.291875
9363.47177
total Q1=
36663.8701
n𝐼𝑛,𝑐𝑜𝑜𝑙𝑖𝑛𝑔
̇ 𝑤𝑎𝑡𝑒𝑟 ,×
̂
H𝐼𝑛 (kJ/s)
-4810.5483
̂ 𝑂𝑢𝑡
n𝑂𝑢𝑡,
̇ ×H
(kJ/s)
36753.83
183.791823
42.6926947
1215.18886
̂ 𝑂𝑢𝑡
n𝑂𝑢𝑡,𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑒𝑑
̇
×H
(kJ/s)
-1531.6333
total Q2=
36663.8701
Since the cooling water inlet flow rate was unknown, then the total molar flow rate of the water exiting
the unit was also unknown. However, using an energy balance where a goal seek method was
possible to use. The goal seek method was used to estimate the molar flow rate of the inlet cooling
water stream. Since n1 is known to be 635.318 mol/s, then n2B can be estimated while also knowing
that n1+n2B=n2+n3.
̂ 1,2B ) was multiplied by the unknown inlet molar flowrate of the cooling water, while H
̂ 1,1 was
(H
multiplied by the known molar flowrate of water in the flue gas (635.318 mol/s). Furthermore, the
̂ 5,2 . Finally, (the total
molar flowrate of the water that condenses (38.751 mol/s) was multiplied by H
outlet molar flowrate of water – the molar flowrate of the water that condensed) was multiplied by
̂ 5,3 . The total inlet kJ/s (Q1) is then subtracted from the total outlet kJ/s (Q2). The goal seek was
H
utilised to guess the value of the inlet molar flowrate of cooling water until Q1 was equal to Q2.
6.2.2 Packed Bed Absorption Column (A-101)
Lead author: Bryn Barker
51
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Stream table for unit A-101
Stream No.
Total flow (kg/s)
Temperature
(°C)
Pressure (bar)
Phase
3A
5
4
12
106.23 265.37 93.28 252.42
40
62.92 40
40
1.2
V
1.1
L
1.2
L
1.1
V
*Note yellow values are calculate in energy balance
Assumptions:
streams 3 and 4 are in thermal equilibrium (Ravi, 2023)
The energy balance for the absorption column was to determine the temperature of stream 5 (the
rich MEA stream), all other streams were pre-defined and had a temperature of 40°C, this was
because 40°C was found to be the optimum temperature for CO2 absorption by MEA. The equation
for the energy balance is shown in equation 6.2.3.1 below.
∑ Hstream 3A + ∑ Hstream 12 = ∑ Hstream 4 + ∑ Hstream 5 + ∆HRXN
eq 6.2.2.1
eq 6.1.3.4
Because the reaction taking place in the absorption column is the absorption of CO2, then the ∆𝐻𝑅𝑋𝑁
term is the enthalpy of absorbtion. The heat of absorption for 30wt% MEA at 40°C with a loading of
0.5mol CO2/ mol MEA was found to be -71.30 Kj/mol CO2 (Kim et al, 2014). This value was multiplied
by the number of moles of CO2 absorbed, 294.32 moles, as calculated in 6.1.3, this gave a value of
-20992 KJ/s. This exothermic reaction is responsible for the rise in temperature of stream 5. It
therefore follows that using a simple 𝑖𝑛 = 𝑜𝑢𝑡 + ∆𝐻𝑅𝑋𝑁 energy balance could be used to determine
the enthalpy of stream 5 and therefore its temperature.
For all streams the enthalpy calculations used a reference temperature of 25°C. The enthalpy
balance calculations are shown in tables 6.2.2.1 to 6.2.2.4 that are located in appendix c. Equation
6.2.2.2 shown below shows the formula used to calculate the enthalpy of stream 5. As in the initial
energy balance the individual component enthalpies of the stream were summed to determine the
total enthalpy of the stream. This process was repeated for streams 3A, 4 and 12.
∑ 𝐻𝑠𝑡𝑟𝑒𝑎𝑚 3𝐴 + ∑ 𝐻𝑠𝑡𝑟𝑒𝑎𝑚 12 − ∑ 𝐻𝑠𝑡𝑟𝑒𝑎𝑚 4 − ∆𝐻𝑅𝑋𝑁 = ∑ 𝐻𝑠𝑡𝑟𝑒𝑎𝑚 5
eq 6.2.2.2
eq 6.1.3.4
From equation 6.2.2.2 the enthalpy of the rich MEA stream was determined to be 34868 KJ/s. Using
goal seek on excel the outlet temperature was determined, this value was determined to be 62.92°C.
This was a 23°C temperature increase.
52
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6.2.3 - E-101:
Lead author: Bryn
Contributing author(s): Ahmed Aldhalei
Stream table for unit E-101
Stream No.
Total flow (kg/s)
Temperature (°C)
Pressure (bar)
Phase
5A
5B
11A 11B
265.37 265.37 252.20 252.20
62.92 102
2.3 1.8
L
L
122
2
L
82.46
1.5
L
*Note yellow values are those calculated by energy balance
Assumptions
Because more rigorous design of the heat exchanger is not required, the heat exchanger is assumed
to be 100% efficient.
The shell and tube heat exchanger exchanges heat between the rich and lean MEA streams,
streams 5 and 11. The allocation of fluids to shell side and tube side depends on multiple parameters
including fouling, corrosiveness of the fluids, what state the fluids are in and the pressure of the
fluids. Because it is cheaper to build thicker tube walls than shell walls. it is advisable to allocate the
higher-pressure stream to the shell side. (ENERQUIP, 2018).Because only the pressure differs
between the 2 process streams the rich MEA stream (stream 5) was set as the shell side fluid.
The temperature of the rich MEA stream was calculated in the energy balance for the packed bed
absorber and the desired temperature of the rich stream entering the stripper was 102°C, as such
the temperature of the tube side was fully defined. The temperature of the lean stream (stream 11)
exiting the stripper was 122°C. Consequently the energy balance calculated the tube side outlet
temperature. The enthalpy balance is shown in equation 6.2.4.1 below.
53
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∆𝐻𝑠𝑡𝑟𝑒𝑎𝑚 5 = ∆𝐻𝑠𝑡𝑟𝑒𝑎𝑚 11
eq 6.2.3.1
eq 6.1.3.4
The Rich stream enthalpy change was found to be 36115KJ this calculation for this was shown in
tables 6.2.4.1 to 6.2.4.2 in appendix c. Using the inlet temperature of the lean stream of 122°C the
goal seek function in excel was used to determine the outlet of the of the heat exchanger, this
temperature was determined to be 82.46°C. The equation used to calculate this is shown below.
36115
𝑛𝑠𝑡𝑟𝑒𝑎𝑚 11 × 𝐶𝑝̇ × (122 − 𝑇)
eq 6.2.3.2
eq 6.1.3.4
6.2.4 Cooler (C-102)
Lead author: Bryn Barker
Stream table for unit C-102
Stream No.
Total flow (kg/s)
Temperature (°C)
Pressure (bar)
Phase
11B
11B
252.20 252.20
82.46 40
1.5
1.2
L
L
cw
cw
542.75 542.75
23
40
3
2.7
L
L
*Note highlighted values are calculated from the energy balance
Assumptions:
the cooler is 100% efficient
There is no absorption in the cooling water stream
For the cooler energy balance the temperatures of all 4 streams were set, the cooling water had a
supply temperature of 23°C and a maximum allowable return temperature of 40°C. For the lean MEA
stream (11B) the temperature inlet was 82.46°C as determined by the energy balance in section
6.2.4. The outlet of the Lean MEA stream was 40°C, this was because the absorption column
required the inlet streams to be at 40°C. Then energy balance carried out determined the mass
flowrate of cooling water required to cool the Lean MEA stream. the enthalpy balance equation is
shown below in equations 6.2.5.1 & 6.2.5.2.
∆𝐻𝑠𝑡𝑟𝑒𝑎𝑚 11𝐵 = ∆𝐻𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟
eq 6.2.4.2
eq 6.1.3.4
54
eq 6.2.4.1
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
∆𝐻𝑠𝑡𝑟𝑒𝑎𝑚 11
= 𝑛𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟
𝐶𝑃 ∆𝑇𝑐𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟
eq 6.2.5.2
eq 6.1.3.4
From these calculations shown in table 6.2.4.2 in appendix C, the cooling duty waseq
calculated
6.2.4.2 to be
38.65MW and the cooling water mass flowrate was determined to be 542.75 Kg/s.
6.2.5 S-101 Energy balance
Lead author: Faisal Kamal
Contributing author(s): Faisal Kamal
•
The assumptions made are the same as those of the mass balance calculations, see chapter
6.7 the stream numbers used and the UBD is the same as well.
• The enthalpies highlighted in the table below were calculated using the equations in this
section, for the sake of space, these enthalpies are included here.
Table 6.2.5.1 Stream table for unit S-101
9
10
11
5B
6
7
8
Stream No.
272.82
20.55
252.27
265.47
28.52
25.32
13.19
Total flow (kg/s)
110.5
120
120
102
102
40
40
Temperature (°C)
2
2
1.9
1.2
1.8
1.7
1.8
Pressure (bar)
11257 43246. 86742.
71994.47
Streams calculated
3581.418
912.72 177.52 2
3
9
Enthalpies (Kw)
L
V
L
L
V
L
V
Phase
The main goal of the energy balance calculation is to determine the duties of the reboiler and
condenser in the system, alongside that, the temperature of stream 10 will also be calculated.
The energy balance will follow simple enthalpy differential calculations to attain the duties of the units
mentioned above, the enthalpies will be calculated in the following manner:
𝐻̇ = 𝑛̇ 𝐶𝑝 ∆𝑇
Equation (6.2.5.1)
Notably, the molar flow rates are used in the equation, the value of the flowrates in table (6.1.4.7) in
the mass balance of the stripper, where used alongside the molar weights in Chapter 3 to convert
the values of the table.
Moving forward, the difference between the enthalpy sum of the inlet streams and the enthalpy sum
of the outlet streams is equal to the duty required in a unit, and therefore:
𝐷𝑢𝑡𝑦 𝑖𝑛 𝑎 𝑢𝑛𝑖𝑡 = ∆𝐻̇ = 𝛴𝐻̇𝑖𝑛 − 𝛴𝐻̇𝑜𝑢𝑡
Equation (6.2.5.2)
6.2.5.1 Condenser Duty
With the understanding of the method used for the calculations, streams 6,7 and 8 are the first
enthalpies to be calculated, and the duty of the condenser is the first to be attained. the specific heat
capacity at the appropriate temperature must be calculated for all 7 streams through. The specific
heat coefficients alongside the calculation alongside the rest of the 𝐶𝑝 calculations will ne be
provided in appendix C. These values were multiplied by the molar flowrate according to equation
(6.2.5.1) to attain the enthalpies. The enthalpy values where then converted to Kj/s. This provides
all the values need to calculate the duty required for the condenser where:
𝐷𝑢𝑡𝑦 𝑜𝑓 𝑡ℎ𝑒 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑜𝑟 = ∆𝐻̇ = 𝐻̇2 − (𝐻̇3 + 𝐻̇4 )
55
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
And so,
𝐷𝑢𝑡𝑦 𝑜𝑓 𝑡ℎ𝑒 𝑐𝑜𝑛𝑑𝑒𝑛𝑠𝑜𝑟 = 3581.41 − (912.72 + 177.52) = 2491.17
𝐾𝑗
𝑜𝑟 2.491 𝑀𝑤
𝑠
The calculation above shows the duty of the condenser is 2.49 Mw, from this calculation, the amount
of cooling water required in the condenser can also be calculated.
The specific heat of water was calculated, see appendix c, the value attained was 1.28 Kj/mol. This
value was divided by the duty requirement to provide the required cooling water flow rate:
𝐾𝑗
𝑠 = 1942.39 𝑚𝑜𝑙 𝑜𝑟 34.98 𝐾𝑔
𝐶𝑜𝑜𝑙𝑖𝑛𝑔 𝑤𝑎𝑡𝑒𝑟 𝑟𝑒𝑞𝑢𝑖𝑟𝑒𝑑:
𝐾𝑗
𝑠
𝑠
1.28
𝑚𝑜𝑙
2491.17
6.2.5.2 Energy balance around the column
At this stage of the calculations, all enthalpies can be attained using equation (6.2.5.1) with the
exception of stream 9. This is because the temperature of stream 9 is unknown, therefore, the
calculating the specific heat capacity, and as an extension, the enthalpy of stream 9 is not possible
in accordance with the method used up to this point. Therefore, to calculate the enthalpy of stream
9, a general energy balance is done around the stripper column itself. In this case, the inlets to the
column would be streams 5B, 7 and 10. The outlets would be streams 6 and 9.
𝐻̇5𝐵 + 𝐻̇7 + 𝐻̇10 = 𝐻̇6 + 𝐻̇9
Equation (6.2.5.3)
And therefore, in order to calculate stream 9, the enthalpies of the rest of the required streams must
be calculated first. The enthalpies of streams 6 and 7 were calculated in the condenser duty section,
and thus, only streams 5B and 10 must be calculated. And so, the enthalpies of both streams were
calculated.
Notably, in the case of stream 10, the boil up stream, the value shown does not consider the heats
of reaction and vaporisation occurring in the reboiler. These values must be added to the enthalpy
of stream 9 before proceeding, in order to calculate these values, the heat of reaction is obtained
from the absorber calculations, as it is known that the absorption of CO2 into aqueous amines is an
exothermic reaction and therefore in order to strip CO2 from amines the required amount of energy
is the same as the amount of energy released during the absorption process (Srisang, 2017).And
therefore, if the heat of absorption is -71.3 Kj/mol of CO2, then the heat of desorption is 71.3 Kj/mol
of CO2. Furthermore, the heat of water vaporisation was obtained from literature to be 40.65 Kj/mol
(Ayala and Rogstad, 2020).
Having acquired the heat of reaction and the heat of vaporisation of water, table (6.2.5.7) from the
stripper mass balance can be used to obtain the calculated amounts of water vaporisation and CO 2
desorption in the reboiler’s boil up stream, stream 10. The table shows 7.96 Kg/s of CO2 desorbed
and 12.12 Kg/s of water vaporised.
56
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
With the amounts of water vaporisation and CO2 reaction known, they are converted to mol/s to
calculate the enthalpy of reaction and vaporisation, Thus, the enthalpies of reaction and water
vaporisation can be calculated as follows:
𝑚𝑜𝑙
𝐾𝑗
∗ 71.3 = 12906.7
𝑠
𝑠
𝑚𝑜𝑙
𝐾𝑗
= 673.20
∗ 40.65 = 27365.89
𝑠
𝑠
𝐻̇𝑟𝑒𝑏 𝑟𝑒𝑐 = 181.02
𝐻̇𝑟𝑒𝑏 𝑣𝑎𝑝
As follows, the enthalpy of stream 9 can now be calculated in addition to the heat of reaction and
vaporisation. This yields a value of ∆𝐻̇6 = 43246.3 Kj/s.
Finally, the enthalpy of stream 9 can now be calculated in accordance with equation (6.2.5.3):
𝐻̇5 = 71994.47 + 912.72 + 43246.3 − 3581.4 = 112572 𝐾𝑗/𝑠
Notably this enthalpy contains the heat of reaction and water vaporisation in the column, so in
removing these values, the calculation of the temperature of the stream is possible. Similarly, to
stream 10, the enthalpies of stream 9’s heat of reaction and condensation are calculated using the
flowrates in the stream attained from the final mass balance table. These calculations result in:
𝑚𝑜𝑙
𝐾𝑗
∗ 71.3 = 30116.3
𝑠
𝑠
𝑚𝑜𝑙
𝐾𝑗
= −18.21
∗ 40.65 = −740.3
𝑠
𝑠
𝐻̇𝑐𝑜𝑙𝑢𝑚𝑛 𝑟𝑒𝑐 = 422.38
𝐻̇𝑐𝑜𝑙𝑢𝑚𝑛 𝑐𝑜𝑛𝑑
And from this, the enthalpy of stream 9 required to calculate the temperature is 83196. Kj/s. Through
goal seeking in excel in the table below, for the sum of the enthalpies to be equal to 83196 Kj/s by
changing the set temperature, a set temperature of 110.6 degrees is attained.
CP values and calculations. STREAM 9 Tref = 25, Tset =110.68 . Phase L
Comp
cp in j/mol
H5 j/s
H5 in Kj/s
MEA
14557.54311
17250197.88
17250.19788
H2O
6449.930359
64327972.14
64327.97214
CO2
3402.891765
1617886.581
1617.886581
Total
83196.0566
6.2.5.3 Reboiler Duty
In order to calculate the reboiler duty, the enthalpies of streams 9,10 and 7 must be known.
Therefore, the enthalpy of stream 11 was calculated in accordance with equations (6.2.6.1-2) to be
𝐻̇11 = 86742.89 Kj/s. With this value, the duty of reboiler would be equal to:
𝑅𝑒𝑏𝑜𝑖𝑙𝑒𝑟 𝐷𝑢𝑡𝑦 = 𝐻̇11 + 𝐻̇10 − 𝐻̇9
And so, the reboiler duty is calculated to be 46793.16 Kj/s or 46.79 Mw. Furthermore, the attained
reboiler duty can be converted to MJ/Kg by dividing reboiler duty to amount of CO2 generated. This
57
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
yields a value of 3.61 MJ/Kg. This value is corporate by several similar studies from literature under
the process conditions similar to those used in this energy balance (Biermann et al., 2022. Mathisen
et al., 2013)
6.2.6 Compressor train (COM-1 to COM-4, C-102 to C-105, D-101&D-102)
Lead author: Abdelrahman Abouelela
Contributing author(s): N/A
Compressor Energy Balance:
Assumptions:
- All Water and MEA is removed by the dryer.
- There is no pressure drop from the coolers.
After the water and MEA is removed from the stream via the adsorption dryer, the pure CO2 stream
enters the compressor train at 40℃ and 1.7 bar and must enter the storage pipeline at 25℃ and 50
bar. The first compressor must compress the stream from 1.7 bar to 14 bar, however, this creates a
temperature increase due to the relationship between pressure and temperature. The outlet
temperature of the stream must be calculated to know the cooling duty of the interstage cooler. The
work done by the compressor must also be calculated. This was done through the use of the
following method:
Firstly, the polytropic efficiency, EP, of the compressor must be found from figure 6.2.7.1. However,
to find the polytropic efficiency from the chart, the compression ratio must first be found. This is
calculated through the use of:
𝑃2
𝑃1
14
= 1.7 = 8.24 (3 sf) (equation 6.2.6.1)
The polytropic efficiency can then be extrapolated from the chart and it is found to be 86% (0.86) for
the first compressor.
After that, an initial value for the polytropic temperature exponent, m, is to be calculated through the
use of:
𝑚=
(𝛾−1)
𝛾𝐸𝑝
(1.29−1)
= 1.29∗0.86 = 0.26 (3 sf)
(equation 6.2.6.2)
Then, the outlet temperature, T2, can be calculated through the use of:
𝑃
𝑚
𝑇2 = 𝑇1 ∗ (𝑃2 )
1
14 0.26
= 313 ∗ (1.7)
= 543K → 270℃ (3 sf) (equation 6.2.6.3)
Since the outlet temperature has been calculated, the mean temperature and pressure are
calculated through the use of:
𝑇𝑚𝑒𝑎𝑛 =
𝑇1 +𝑇2
2
(equation 6.2.6.4)
58
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
𝑃1 +𝑃2
2
𝑃𝑚𝑒𝑎𝑛 =
For the first compressor, 𝑇𝑚𝑒𝑎𝑛 =
(equation 6.2.6.5)
318 + 543
2
= 431 K and 𝑃𝑚𝑒𝑎𝑛 =
1.7+14
2
= 7.9 bar
The mean temperature and pressure are then used to find the reduced temperature, Tr, and reduced
pressure, Pr.
𝑇𝑟 =
𝑇𝑚𝑒𝑎𝑛
𝑇𝑐
=
431
304
= 1.42 (3 sf)
(equation 6.2.6.6)
𝑃𝑟 =
𝑃𝑚𝑒𝑎𝑛
𝑃𝑐
= 73.8 = 0.11 (3 sf)
(equation 6.2.6.7)
7.9
The reduced temperature and pressure values are then used to find the compressibility factor, Z,
and compressibility functions X and Y, from their respective charts.
From figures 6.2.7.2, 6.2.7.3 and 6.2.7.4 respectively, Z = 1, X = 0.06, Y = 1.02, for the first
compressor.
After that, a revised value for the polytropic temperature exponent, m, is to be calculated through the
use of:
𝑚=
𝑍𝑅 1
(
𝐶𝑃 𝐸𝑃
+ 𝑋) =
1 ∗ 8.314
1
(
+
37.35
0.86
0.06) = 0.27 (3 sf)
(equation 6.2.6.8)
Then, the polytropic exponent, n, is to be calculated through the use of:
1
1
𝑛 = 𝛾−𝑚(1+𝑋) = 1.29−0.27(1+0.06) = 1.37 (3 sf)
(equation 6.2.6.9)
Using the values obtained, the work done by the first compressor can be calculated through the use
of:
𝑛−1
𝑊=𝑍
𝑅𝑇1 𝑛
𝑃
𝑛
[( 2 )
𝑀 𝑛−1 𝑃1
1.37−1
− 1] = 1
8.314 ∗ 318 1.37
1.8 1.37
[( )
44
1.37−1 1.7
− 1] = 348 j/mol (3 sf) (equation 6.2.6.10)
The actual work done is calculated through the formula:
𝑊𝑜𝑟𝑘 𝐷𝑜𝑛𝑒
348
𝐴𝑐𝑡𝑢𝑎𝑙 𝑊𝑜𝑟𝑘 𝐷𝑜𝑛𝑒 (𝑊𝐷0 ) = 𝑃𝑜𝑙𝑦𝑡𝑟𝑜𝑝𝑖𝑐 𝐸𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦 = 0.86 = 405 j/mol (3 sf)
(equation 6.2.6.11)
To convert the work done from j/mol to kW, the flow rate of CO2, in mol/s, must be converted to kg/s.
This is done through the use of:
𝜐=
𝜐0 ∗ 𝑀𝑟
1000
=
294.42 ∗ 44
1000
= 13 kg/s
59
(equation 6.2.6.12)
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
The flowrate in kg/s must then be multiplied by the work done in j/mol and divided by 1000 to obtain
the work done on a kW basis:
𝑊𝐷 =
𝜐 ∗ 𝑊𝐷0
1000
=
13 ∗ 403
1000
= 119 kW (3 sf)
(equation 6.2.6.13)
This process is repeated for the next three compressors, with the calculated outlet temperature and
pressure conditions of the CO2 stream from the cooler being the inlet conditions for the next
compressor.
Cooler EB:
To find the outlet temperature of the stream after it has exited the cooler requires equating and
rearranging the following equation for work done by the cooler:
𝑄 = 𝑚̇𝐶𝑃 (𝑇𝑖𝑛 − 𝑇𝑜𝑢𝑡 )
(equation 6.2.6.14)
𝑚̇𝐶𝑂2 𝐶𝑃 𝐶𝑂2 (𝑇𝑖𝑛 − 𝑇𝑜𝑢𝑡 )𝐶𝑂2 = 𝑚̇𝑤𝑎𝑡𝑒𝑟 𝐶𝑃 𝑤𝑎𝑡𝑒𝑟 (𝑇𝑖𝑛 − 𝑇𝑜𝑢𝑡 )𝑤𝑎𝑡𝑒𝑟
𝑇𝑜𝑢𝑡 𝐶𝑂2 = [
𝑚̇𝑤𝑎𝑡𝑒𝑟 𝐶𝑃 𝑤𝑎𝑡𝑒𝑟 (𝑇𝑖𝑛 −𝑇𝑜𝑢𝑡 )𝑤𝑎𝑡𝑒𝑟
𝑚̇𝐶𝑂2 𝐶𝑃 𝐶𝑂2
(equation 6.2.6.15)
] − 𝑇𝑖𝑛 𝐶𝑂2(equation 6.2.6.16)
For the first cooler, an initial value for the flow rate of water was chosen (15 kg/s). Furthermore, the
cooling water has a summer temperature of 23℃ and a maximum return temperature of 40℃. This
gives an outlet stream temperature of:
𝑇𝑜𝑢𝑡 𝐶𝑂2 = [
15 ∗ 4180 ∗ (313−296)
]−
12.95 ∗ 849
544 = 447K → 174℃
To find the work done by the cooler, equation 6.2.6.14 is utilised:
𝑄 = 15 ∗ 4180 ∗ (313 − 296) = 1066 kW
Since the equations are equated, the work done on the water is equal to the work done on the CO2
stream, meaning the work done of 1066 kW is doubled to 2132 kW, giving the total work done by
the first compressor.
This process is repeated for the next two coolers, with the calculated outlet temperature and pressure
conditions of the CO2 stream from the compressor being the inlet conditions for the next cooler.
However, for the final cooler, the unknown variable switches from the outlet temperature of the CO2
60
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
stream to the flow rate of water required so equation 6.2.7.15 is rearranged to calculate the unknown
flow rate required:
𝑚̇𝑤𝑎𝑡𝑒𝑟 =
𝑚̇𝐶𝑂2 𝐶𝑃 𝐶𝑂2 (𝑇𝑖𝑛 −𝑇𝑜𝑢𝑡 )𝐶𝑂2
𝐶𝑃 𝑤𝑎𝑡𝑒𝑟 (𝑇𝑖𝑛 −𝑇𝑜𝑢𝑡 )𝑤𝑎𝑡𝑒𝑟
(equation 6.2.6.16)
For the final cooler, the flow rate of water is:
𝑚̇𝑤𝑎𝑡𝑒𝑟 =
12.95 ∗ 849 ∗ (403−298)
4180 ∗ (313−296)
= 16.3 kg/s (3 sf)
Table 6.2.6.1: Stream table for compressors in compressor train
Unit
COM-101
COM-102
COM-103
In/Out
In
Out
In
Out
In
Out
In
Out
Stream No.
13
14
15
16
17
18
19
20
CO2
(kg/s)
12.95
12.95
12.95
12.95
12.95
12.95
12.95
12.95
Temperature
(°C)
40
270
173
253
155
199
102
130
Pressure
(bar)
1.7
14
14
26
26
38
38
50
Work
(kW)
119
flow
Done
Phase
V
91
L
L
COM-104
74
L
L
73
L
L
L
Table 6.2.6.2: Stream table for coolers in compressor train
Unit
C-102
In/Out
In
Out
In
Out
In
Out
In
Out
Stream No.
14
15
16
17
18
19
20
21
CO2
(kg/s)
flow
12.95
12.95
12.95
12.95
12.95
12.95
12.95
12.95
H2O
(kg/s)
flow
15
15
15
15
15
15
16.4
16.4
Temperature
(°C)
270
173
253
155
199
102
130
25
Pressure
(bar)
14
14
26
26
38
38
50
50
Work
(kW)
2132
Phase
Done
L
C-103
C-104
2132
L
L
C-105
2132
L
L
61
2316
L
L
L
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
*Highlighted values are calculated
6.3 Process Integration
6.3.1 Hot and cold utility requirements
Table 6.3.1.1: hot and cold plant utility requirements
Description
C-101
C-102
C-103
C-104
C-105
REB
Hot Utility (KW)
Cold utility (KW)
38500
1065.9
1065.9
1065.9
1158.0
46793
6.3.2 Selection of streams for PI
4 streams were chosen for the pinch point analysis, these being the rich MEA stream, the lean MEA
stream the reboiler stream and the condenser stream. The streams between compressors were not
selected as the cost of building pipelines to the start of the process would likely be not feasible. For
similar reasons the flue gas outlet stream from the power plant was not selected for the pinch
analysis as it would be impractical to pump it far downstream of the absorber only to have to pump
it back to the absorber. Of the 4 streams selected, the lean MEA and condenser were both hot
streams, while the rich MEA and reboiler streams were both cool streams.
6.3.3 Minimum approach temperature
A minimum approach temperature of 20°C was selected for the pinch point analysis. Typical minium
approach temperatures range from 5°C to 30°C. In general a lower minimum approach temperature
decreases the requirement of additional utilities however this comes at the cost of the present utilities
needing to be larger in size, hence increasing their capital cost (Towler and Sinnott, 2013). A
minimum approach temperature of 20°C would balance these requirements and be a suitable trade
off.
62
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6.3.4 Application of PI and design of HEN
The following diagrams are provided by: (Xu, 2006)
Figure 6.3.4.1: input data for pinch point analysis
Figure 6.3.4.2: cascade table produced
Figure 6.3.4.3: pinch point diagram
From the pinch point calculations carried out on the provided spreadsheet using the input data shown
in figure 6.3.4.1, the pinch point was determined to be 83°C. From the pinch point diagram shown in
figure 6.3.4.3 it was shown that the only streams capable of exchanging heat were the rich and lean
63
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
MEA streams. Because the Reboiler stream doesn’t come close to the other streams in temperature
additional hot utility is required for it to function. This was provided by the kettle reboiler in the current
plant design. Because the lean MEA stream can only transfer energy to the rich stream above the
pinch as shown by figure 6.3.4.3, the lean stream will require additional cooling utility. This was
already incorporated into the energy balance via the addition of a cooler after the rich-lean heat
exchanger. Since cooling of the condensate stream typically takes place in the condenser it would
require considerable redesign to incorporate the condensate stream into a heat exchanger network.
The pinch point analysis indicates that the current design set out in the report is optimised and no
further redesign is required.
64
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6.4 Final Process Flow Diagr
Equiptment code
Equiptment Name
Design Temp. (°C)
Design Pressure
(bar)
DCC-101
Direct
contact
cooler
140
A-101
Packed
bed
absorber
40
1.2
1.2
E-101
Shell and
tube heat
exchanger
63 Rich
122 Lean
2.3 Rich
2.0 Lean
C-102
Cooler
S-101
Stripping
unit
COM 1 to 4
Compressor
train
82.5
102
1.5
1.8
270, 253,
199, 130
14, 26 ,38,
50
Utility
Cooling
water
Equiptment C-102
23, 3
Tin (°C) pin
(bar)
Tout (°C)
40, 2.7
Pout (bar)
Flow (kg/s) 542.75
Stream NO.
Component
flow (kg/s)
MEA
CO2
H2O
N2
O2
Total flow
(kg/s)
Enthalpy (kW)
Temperature
(°C)
Pressure (bar)
Phase
1
2B
2
3
3A
5
12
4
4A
WWin
WWout
5A
5B
11A
11B
0
14.39
11.44
78.011
3.091
0
0
574.203
0
0
0
0.089
574.90
0
0
0
14.305
10.74
78.011
3.091
0
14.39
10.74
78.01
3.09
0
25.91
167.62
0
0
0
12.95
167.62
0
0
0
1.44
10.74
78.01
3.09
0.00016
1.44
10.74
78.01
3.09
0
0
1.35
0
0
0
0
1.35
0
0
71.84
25.91
167.62
0
0
71.84
25.91
167.62
0
0
71.84
12.95
167.40
0
0
71.84
12.95
167.40
0
0
106.93
574.99
574.99
106.23
106.23
265.37
252.42
93.37
93.28
1.35
1.36
265.37
265.37
252.20
252.20
41474.4
2
-4810.55
34548.6
2
2114.12
1747
34869
13706
1578
1578
1578
1578
71994.
47
71994.4
7
86742
38649
140
23
40
40
40
63
40
40
40
23
23
63
102
122
82.5
1.2
V
3
L
1.1
L
2.9
V
1.2
V
1.1
L
1.2
L
1.1
V
40
1.1
3
L
1.9
L
2.3
L
1.8
L
2
L
1.5
L
65
Stream NO.
Component
flow (kg/s)
MEA
CO2
H2O
N2
O2
Total flow
(kg/s)
Enthalpy
(kW)
Temperature
(°C)
Pressure (bar)
Phase
Cooling
water
C-103
Cooling
water
C-104
Cooling
water
C-105
Cooling
water
C-106
23, 3
23, 3
23, 3
23, 3
P-101
63,1.1
P-102
122, 1.7
40, 2.7
40, 2.7
40, 2.7
40, 2.7
63, 2.3
2.0
15
15
15
16.30
265.37
265.18
Pumping power
6
7
8
9
10
11
13
14
15
16
17
18
19
20
21
0.17
26.56
11.80
0
0
11.58
13.60
0.15
0
0
0.016
12.96
0.22
0
0
72.28
20.92
178.62
0.46
7.97
12.12
71.82
12.96
167.50
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
0
12.96
0
0
0
38.52
25.33
13.19
272.83
20.55
252.28
12.96
12.96
12.66
12.96
12.96
12.96
12.96
12.96
12.96
3581.4
1
912.72
177.5
112572
43246.3
1
86742
177.5
119
2132
91
2132
74
2132
73
2316
102.78
40
40
110.6
120
120
40
270
173
252
155
199
102
120
25
1.8
V
1.8
L
1.8
V
2
L
2
V
2
L
1.7
V
14
L
14
L
26
L
26
L
38
L
38
L
50
L
50
L
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6.5 Auxiliary unit power requirement calculations
Pump calculations:
P-102 increased the pressure of the lean MEA stream, (stream 5) from 1.1 bar at the exit of the
absorption column to 2.3 bar at the exit of the pump. This pressure increase was decided based on
the upstream stripper pressure requirement of 1.8 bar. The equations used to calculate pump power
are shown below:
𝑞𝜌𝑔ℎ
(3 × 106 ) Equation 6.5.1 Pump power calculation
(EngineeringToolbox, 2003)
The head, denoted h in equation 6.4.1 would typically be calculated more rigorously in a pump design
however typical head ranges are given in literature, for centrifugal pumps the typical range of head
is between 10-50m. (Coulson & Richardson, 2005). The upper limit was selected as to give an upper
estimate to the required pump power.
the average fluid density was determined using the sum of the individual component densities. The
densities of fluids used are shown in table 6.5.1 below.
Table 6.5.1: fluid densities, (EngineeringToolbox, 2003) (cameochemicals, 1999)
Component
ρ at 63°C
ρ at 120°C
MEA
979.29
1181
H2O
981.61
943.08
CO2
598
990.36
determine the volume flowrate in
𝑚3
ℎ𝑟
to
the formulas below were used. The stream 5A volume flowrates
are displayed in table 6.5.2:
𝑀̇
𝑚3
× 3600 = 𝑣𝑜𝑙𝑢𝑚𝑒𝑡𝑟𝑖𝑐 𝑓𝑙𝑜𝑤𝑟𝑎𝑡𝑒
𝜌
ℎ𝑟
71.84
𝑚3
× 3600 = 264.11
979.29
ℎ𝑟
Equation 6.5.2
Equation 6.5.3
Table 6.5.2: volume flowrates of stream 5A
component mol FR
mass
𝐾𝑔
𝑠
MEA
H2O
CO2
1177.682 71.84
9312.413 167.62
588.8411 25.91
FR/ Vol FR /
𝑚3
ℎ𝑟
264.11
614.75
155.97
The sum of the individual volume flowrates shown in table 6.5.2 was found to be 1034.83
𝑚3
ℎ𝑟
using
equation 6.4.1 where 𝜌 was the the average density of the fluid the pump power can be found with
66
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
𝐾𝑔
the following equation. for stream 5A this density was calculated to be 852.94𝑚3. The calculation for
this is shown in equation 6.4.4 below
1034.83 × 852.94 × 9.81 × 50
= 120.26𝐾𝑤
(3.6 × 106 )
Equation 6.5.4
Using the same calculations, the power needed for pump 102 was calculated at its specific density
and flowrate (Appendix C). The power of P-102 was found to be 123.2002 Kw, the pump raises the
pressure of the stream to 2 bar in order to take the pressure drop of the heat exchanger and the
cooler into account and deliver the stream at 1.2 bar, the optimum pressure in the absorber.
7. CONCLUSIONS
A post-combustion carbon capture method has been decided as the most viable process for this
project along with MEA as the amine solution due to its increased absorption capabilities stemming
from its high reactivity. Raw materials used for construction should be bought from Chinese suppliers
to benefit economically from cheaper prices. Likewise, coal bed methane should also be supplied
from China for the aforementioned reason. On the other hand, MEA should be bought from a local
supplier due to its stable market price and for the flexibility from the supplier to FAABS. Some of the
chemicals that are being used in operation can be very dangerous if handled improperly, leading to
the implementation of exposure controls, proper storage and disposal guidelines, such as in the case
of wastewater where it is treated before disposal to prevent any negative environmental impacts.
The units in the operational process have all achieved their desired benchmarks, such as the
absorber column absorbing over 90% of the CO2 available from the flue gas and the stripper
producing a CO2 stream with a purity over 95 wt%. Once the stream has been compressed and
liquified, it was sent to the underground pipeline to be stored. Calculations illustrate that ≈13 kg/s of
CO2 will be stored when the plant is fully operational. FAABS opted to store the CO2 rather than sell
it to gain Australian Carbon Credit Units (ACCUs). As this project is projected to capture and store
370,000 tons of CO2 per year, worth more than $4.5 million, this makes the project economically
viable while simultaneously positively impacting the environment by reducing CO2 emissions.
67
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
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9. Appendices
APPENDIX A: PROJECT PLAN
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APPENDIX B - Raw Materials and Products appendices:
Table B.1: potential steel & concrete suppliers
Local Suppliers:
Blue Scope Steel Limited
Sims Limited
Bisalloy Steel Limited
Equatorial Resources Limited
Local Suppliers:
Hanson Australia
Cement Australia
Holcim Australia
Brickworks
Steel
International Suppliers:
Park side steel (UK)
Nucor corporation (USA)
Ezz Steel (Egypt)
China Baowu Steel Group (China)
Concrete
International Suppliers:
Foundation Developments (UK)
Baker Construction Enterprises Inc. (USA)
China National Building Material Group (China)
Lafarge Holcim (Switzerland)
Appendix B.2 : Potential MEA & coal bed methane suppliers
Local Suppliers:
MEA
International Suppliers:
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Brenntag
CSBP
Redox Pty Ltd
Orica
Local Suppliers:
Woodside Petroleum
Santos Ltd
Origin Energy
AMPOL
Dow Chemical Company (USA)
SABIC (Saudi Arabia)
BASF (Germany)
INEOS (UK)
Coal bed methane
International Suppliers:
Shell (Netherlands)
British Petroleum (UK)
ExxonMobil (USA)
China United coalbed Methane Corp. Ltd. (China)
Appendix B.3 –: Market price of MEA in different geographical regions (businessanalytiq, 2023)
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Appendix B.4 : Projections for future global CO2 demand (IEA, 2019)
APPENDIX C - Mass and Energy Balance appendices
MSDS:
O2: Air Liquide, 2021. O2 MSDS. Air Liquide Security Filtering. Available from:
http://docs.airliquide.com.au/msdsau/AL605.pdf.
N2: Honeywell. 2017. N2 MSDS. Honeywell Burdick & Johnson. Available from:
https://shop.chemsupply.com.au/documents/LC365M.pdf
MEA: CHEM supply Australia. 2020. MEA MSDS.CHEM-SUPPLY PTY LTD. Available from:
https://shop.chemsupply.com.au/documents/ML0271CHHZ.pdf
CO2: Air Liquide. 2021. CO2 MSDS. Air Liquide Security Filtering. Available from:
http://docs.airliquide.com.au/msdsau/AL062.pdf.
H2O: LAB CHEM. 2020. H2O MSDS. LABCHEM, INC. Available from:
https://www.labchem.com/tools/msds/msds/LC26750.pdf
Exposure limits:
Australian Government. N.D. Exposure Standard Documentation. Available from:
https://hcis.safeworkaustralia.gov.au/ExposureStandards/Document?exposureStandardID=109.
Table 6.2: Table of Cp values used in energy balance, (henni et al 2007), (Coker, 2007), (Felder & Rousseau, 2018)
Source
ACS
publication
Felder
Ludwig’s
Ludwig’s
Ludwig’s
Ludwig’s
Componant
MEA (L)
CP (J/mol.K)
a
b
170
H2O (L)
H2O (g)
CO2 (g)
N2 (g)
O2 (g)
75.4
33.993
27.437
29.342
29.526
-8.42*10^-3
0.042315
-3.54*10^-3
-8.*10^-3
C
d
2.9906*10^-5
1.9555*10^-5
1.0076*10^-5
3.8083*10^-5
-1.7825*10^-8
3.9968*10^-9
-4.312*10^-9
-3.2629*10^-8
Absorbtion column appendices:
Table 6.2.2.1: stream 3A Cp enthalpy table
Stream (3A) Flue gas from
CP (J/mol.K)
DCC
CO2 (g)
579.7923
H2O (g)
504.6641
O2 (g)
441.46798
N2 (g)
436.17389
Sum of individual componant enthalpies (KJ)
H (KJ)
188.4953
301.0661
42.63664
1215.219
1747.417
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Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Table 6.2.2.2: stream 4 Cp enthalpy table
Stream (4) cleaned flue gas to
CP (J/mol.K)
atmosphere
CO2 (g)
579.7923
H2O (g)
504.6641
O2 (g)
441.46798
N2 (g)
436.17389
Sum of individual componant enthalpies (KJ)
H (KJ)
18.96697
301.0661
42.63664
1215.219
1577.8889
Table 6.2.2.3: stream 12 Cp enthalpy table
Stream (x) L-MEA
CP (J/mol.K)
MEA (l)
2550
H2O (l)
1131
CO2 (aq)
579.7923
Sum of individual componant enthalpies (KJ)
H (KJ)
3003.089
10532.34
170.7028
13706.13
Stream (x) R-MEA
CP (J/mol.K)
MEA (l)
6446.06
H2O (l)
2859.0172
CO2 (aq)
1109.38197
Sum of individual componant enthalpies (KJ)
H (KJ)
7591.4098
26624.348
653.24968
34869
Heat exchanger:
Table 6.2.3.1: stream 11 Cp enthalpy table
Stream (11A-B) L-MEA
MEA (l)
H2O (l)
CO2 (aq)
CP (J/mol.K)
H (KJ)
6630
7915.539025
27724.85149
474.59
36115
2904.6
1567.264
Sum of individual componant enthalpies (KJ)
Table 6.2.3.2: stream 5 Cp enthalpy table
Stream (5A-5B) R-MEA
MEA (l)
H2O (l)
CO2 (aq)
CP (J/mol.K)
6630
2940.6
1567.264
Sum of individual componant enthalpies (KJ)
H (KJ)
7808.033
27384.08
922.8692
36115
Cooler exchanger:
Table 6,.2.4.1: stream 11B Cp enthalpy balance
Stream (11B) L-MEA
MEA (l)
H2O (l)
CP (J/mol.K)
7218.2
1271.79
H (KJ)
8500.74532
374.441873
9
Table 6.2.3.4: stream 5 Cp
enthalpy table
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
CO2 (aq)
3201.48
Sum of individual componant enthalpies (KJ)
29774.5865
38650
Table 6.2.4.2: stream cw Cp enthalpy balance
Stream (x) cool water C-102
H2O (l)
CP (J/mol.K)
1281.8
H (KJ)
38650
Stripper energy balance:
Comp
MEA
H2O
CO2
Total
A
Comp
MEA
H2O
CO2
Total
A
Comp
MEA
H2O
CO2
Total
A
Comp
MEA
H2O
CO2
Comp
MEA
H2O
CO2
CP values and calculations. STREAM 8 Tref = 25, Tset =40. Phase V
C
D
cp in j/mol
H4 j/s
0
0
0
2550
663.6592337
-0.008419
0.000029906
-1.78E-08
5.05E+02
6148.08761
0.042315
-0.000019555 3.9968E-09
579.8197
170710.8221
B
170
33.933
27.437
CP values and calculations. STREAM 7 Tref = 25, Tset =40. Phase L
C
D
cp in j/mol
H3 j/s
0
0
0
2550
6331.243302
-0.039953
-0.00021103
5.35E-07
1131.33146
727233.1897
0.042315
-0.000019555 3.9968E-09 579.8196788
179160.26
B
170
92.053
27.437
CP values and calculations. STREAM 6 Tref = 25, Tset =102.78. Phase V
C
D
cp in j/mol
H2 j/s
0
0
0
13090
35907.16635
-0.0084186
0.000029906
-1.78E-08 2603.352232
1705183.542
0.042315
-0.000019555 3.9968E-09 3049.861058
1840327.653
B
170
33.933
27.437
CP values and calcuations lean mea
B
C
D
cp in j/mol H2 j/s
H2 in kj
170
0
0
0
16150 19015363.6 19015.36
92.053 -0.039953
-0.00021103
5.35E-07 7.16E+03 66612143.6 66612.14
27.437 0.042315
-0.000019555 3.9968E-09 3788.424 1115389.7 1115.39
A
A
B
170
92.053
27.437
C
0
-0.039953
0.042315
Comp
MEA
H2O
CO2
A
B
170
33.933
27.437
H4 in Kj/s
0.66365923
6.14808761
170.710822
177.522569
H3 in Kj/s
6.331243302
727.2331897
179.16026
912.724693
H2 in Kj/s
35.90716635
1705.183542
1840.327653
3581.418361
CP values and calcuations boil up
C
D
cp in j/mol
H3 in j/s
H3 in kj
0
0
0
16150 121841.9966 121.842
-0.0084186 0.000029906
-1.78E-08
3.22E+03
2.17E+06 2165.77
0.042315 -0.000019555 3.9968E-09 3788.423549 685795.8921 685.7959
CP values and calcuations TO REB
D
cp in j/mol
H2 j/mol
0
0 14557.54311 17250197.88
-0.00021103
5.35E-07 6449.930359 64327972.14
-0.000019555
3.9968E-09 3402.891765 1617886.581
Compressor train:
Figure 6.2.6.1: Typical efficiencies for reciprocating compressors (Sinnott, 2005)
9
H2 in kj
17250.19788
64327.97214
1617.886581
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Figure 6.2.6.2: Compressibility factors of gases and vapours (Sinnott, 2005)
Figure 6.2.6.3: Generalised compressibility function X (Sinnott, 2005)
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Figure 6.2.6.4: Generalised compressibility function Y (Sinnott, 2005)
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Table 6.2.6.3: Known thermophysical properties of CO2 (engineeringtoolbox, 2018):
CP (j/mol)
37.35
Specific heat ratio (ɣ)
1.29
Critical Temperature (K)
304
Critical Pressure (bar)
73.8
Table 6.2.6.4: CP values (engineeringtoolbox, 2018):
CP of CO2 (j/kg K)
4180
CP of H2O (j/kg K)
849
PUMP:
P-103
1.2 bar req'd
comp
1.7 bar out of strp MEA
H2O
CO2
mol FR
MW
1177.422
9300.229
294.4205
mass FR KG/s
Vol FR m^3/s Vol FR m^3/hr
61 71.82274 0.060815187 218.934673
18 167.4041 0.177507871 639.028335
44 12.9545 0.015294573 55.0604643
tot vol FR m^3/hr 913.023472
Centrifugal selected
tot vol FR
857.963 m^3/hr
Hysys value ?
m^3/hr
Pump power123.2002
(Kw)
Pump power (MW)
0.1232
APPENDIX D - MINUTES OF MEETINGS WITH SUPERVISOR:
Appendix D.1 - Week 1 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 – GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 am on 03/10/2023.
1. Present: Sari Miari, Abdelrahman Abouelela, Ahmad Aldhalei, Bryn Barker, Faisal Kamal
2. Apologies for absence: N/A
3. Absentees: N/A
4. The meeting was chaired, and minutes were taken by Sari Miari
5. Minutes of the previous meeting of the group held on: N/A
6. Matters arising from the previous meeting: Discussed the things we need to research about (post-com,
pre-comb, oxyfuel-comb and types of solvents used to absorb CO2)
7. Minutes of items arising from the meeting agenda: Decided who will research about each section
8. Any other business: N/A
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
9. Chairperson for the next meeting: Sari Miari
10. Next meeting to be held at 11:00 am on 10/10/2023 in Prof Tim Cockerill office.
Appendix D.2 - Week 2 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 – GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 am on 10/10/2023.
1. Present: Sari Miari, Abdelrahman Abouelela, Ahmad Aldhalei, Bryn Barker, Faisal Kamal
2. Apologies for absence: N/A
3. Absentees: N/A
4. The meeting was chaired, and minutes were taken by Sari Miari
5. Minutes of the previous meeting of the group held on: 03/10/2023
6. Matters arising from the previous meeting: Discussed the gantt chart, and picked the process route
which was post-combustion and chose monoethanolamine (MEA) as our solvent for absorption
7. Minutes of items arising from the meeting agenda: Chose who was going to write about each advantage
and disadvantage of each solvent and process discussed, also chose who to start the mass balance
calculations (which is shown in the gantt chart)
8. Any other business: N/A
9. Chairperson for the next meeting: Sari Miari
10. Next meeting to be held at 11:00 am on 10/10/2023 in Prof Tim Cockerill office.
Appendix D.3 - Week 3 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 – GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 am on 17/10/2023.
1. Present: Sari Miari, Abdelrahman Abouelela, Ahmad Aldhalei, Bryn Barker, Faisal Kamal
2. Apologies for absence: N/A
3. Absentees: N/A
4. The meeting was chaired, and minutes were taken by Sari Miari
5. Minutes of the previous meeting of the group held on: 03/10/2023
6. Matters arising from the previous meeting: Discussed the gantt chart, and picked the process route
which was post-combustion and also chose monoethanolamine (MEA) as our solvent for absorption
7. Minutes of items arising from the meeting agenda: Chose who was going to write about each advantage
and disadvantage of each solvent and process discussed, also chose who to start the mass balance
calculations (which is shown in the gantt chart)
8. Any other business: N/A
9. Chairperson for the next meeting: Sari Miari
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
10. Next meeting to be held at 11:00 am on 10/10/2023 in Prof Tim Cockerill office.
Appendix D.4 - Week 4 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 24/10
1. Present:- Abdelrahman Abouelela, Sari Miari, Bryn Barker, Faisal Kamal, Ahmed Aldhalei
2. Apologies for absence:- N/A
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 17/10
Chose who was going to write about each advantage and disadvantage of each solvent and process discussed,
also chose who to start the mass balance calculations (which is shown in the gantt chart)
6. Matters arising from the previous meeting:
- Advantages and disadvantages of solvents to be researched
7. Minutes of items arising from the meeting agenda
- Continue with Mass and Energy Balance (everyone involved)
- Meet on Wednesday 25/10 to decide the individual equipment design
- References to be sorted to Leeds Harvard
8. Any other business N/A
9. Chairperson for the next meeting: Abdelrahman Abouelela
10. Next meeting to be held at 11:00 on 31/10 in SCAPE 3.06
Appendix D.5 - Week 5 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 31/10
1. Present:- Abdelrahman Abouelela, Sari Miari, Bryn Barker, Faisal Kamal
2. Apologies for absence:- Ahmed Aldhalei
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
5. Minutes of the previous meeting of the group held on 24/10
- Continue with Mass and Energy Balance (everyone involved)
- Meet on Wednesday 25/10 to decide the individual equipment design
- References to be sorted to Leeds Harvard
6. Matters arising from the previous meeting:
- Mass and Energy Balance is the priority for the next three weeks
7. Minutes of items arising from the meeting agenda
- Determine flowrate of solvent
- Continue to work on mass and energy balance (everyone involved)
- Sari works on the direct contact cooler
- Research low NOx burners
8. Any other business N/A
9. Chairperson for the next meeting: Ahmed Aldhalei
10. Next meeting to be held at 11:00 on 7/11 in Energy G.1
Appendix D.6 - Week 6 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 7/11
1. Present:- Abdelrahman Abouelela, Bryn Barker, Faisal Kamal, Sari Miari
2. Apologies for absence:- Ahmed Aldhalei
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 31/10
- Determine flowrate of solvent
- Continue to work on mass and energy balance (everyone involved)
- Sari works on the direct contact cooler
- Research low NOx burners
6. Matters arising from the previous meeting:
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
- Mass and Energy balance is the priority for the next two weeks
7. Minutes of items arising from the meeting agenda
- Faisal and Bryn work together on absorber and separator M&EB
- Abdelrahman starts preliminary work on compressor train M&EB
8. Any other business N/A
9. Chairperson for the next meeting: Ahmed Aldhalei
10. Next meeting to be held at 11:00 on 14/11 in Energy G.13
Appendix D.7 - Week 7 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 14/11
1. Present:- Abdelrahman Abouelela, Bryn Barker, Faisal Kamal, Sari Miari
2. Apologies for absence:- Ahmed Aldhalei
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 7/11
- Faisal and Bryn work together on absorber and separator M&EB
- Abdelrahman starts preliminary work on compressor train M&EB
6. Matters arising from the previous meeting:
- Mass and Energy balance is the priority for the next week
7. Minutes of items arising from the meeting agenda:
- Faisal and Bryn finish absorber and separator M&EB
- Abdelrahman finishes mass and energy balance for compressor train
8. Any other business N/A
9. Chairperson for the next meeting: Ahmed Aldhalei
10. Next meeting to be held at 11:00 on 21/11 in Energy G.13
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
Appendix D.8 - Week 8 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 21/11
1. Present:- Abdelrahman Abouelela, Bryn Barker
2. Apologies for absence:- Ahmed Aldhalei, Sari Miari, Faisal Kamal
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 14/10
- Faisal and Bryn finish absorber and separator M&EB
- Abdelrahman finishes mass and energy balance for compressor train
6. Matters arising from the previous meeting:
- Mass and Energy balance is the priority for the next week
7. Minutes of items arising from the meeting agenda
Start write up for mass and energy balance and any uncompleted chapters.
8. Any other business N/A
9. Chairperson for the next meeting: Bryn Barker
10. Next meeting to be held at 11:00 on 28/11 in Energy G.13
Appendix D.9 - Week 9 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 28/11
1. Present:- Abdelrahman Abouelela, Bryn Barker, Faisal Kamal, Sari Miari
2. Apologies for absence:- Ahmed Aldhalei
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 21/11
- Start write up for mass and energy balance and any uncompleted chapters.
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
6. Matters arising from the previous meeting: N/A
7. Minutes of items arising from the meeting agenda
- Continue write up of mass and energy balance and uncompleted chapters (chapters 3 & 5).
8. Any other business N/A
9. Chairperson for the next meeting: Bryn Barker
10. Next meeting to be held at 11:00 on 5/12 in Energy G.13
Appendix D.10 - Week 10 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 5/12
1. Present:- Abdelrahman Abouelela, Bryn Barker, Faisal Kamal, Sari Miari
2. Apologies for absence:- Ahmed Aldhalei
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 28/11
- Continue write up of mass and energy balance and uncompleted chapters (chapters 3 & 5)
6. Matters arising from the previous meeting: N/A
7. Minutes of items arising from the meeting agenda
- Continue write up of mass and energy balance and chapter 5.
8. Any other business N/A
9. Chairperson for the next meeting: Faisal Kamal
10. Next meeting to be held at 11:00 on 12/12 in Energy G.13
Appendix D.11 - Week 11 meeting minutes:
CAPE3000 DESIGN PROJECT, 2023-24 - GROUP CEE2
Minutes of the Semester-1 Group Meeting with Project Supervisor held at 11:00 on 12/12
1. Present:- Abdelrahman Abouelela, Bryn Barker, Sari Miari
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
2. Apologies for absence:- Ahmed Aldhalei, Faisal Kamal
3. Absentees:- N/A
4. The meeting was chaired and minutes were taken by Abdelrahman Abouelela
5. Minutes of the previous meeting of the group held on 5/12
- Continue write up of mass and energy balance and chapter 5.
6. Matters arising from the previous meeting: N/A
7. Minutes of items arising from the meeting agenda
Finish final touches on individual write-ups and collate work to format.
8. Any other business N/A
9. Chairperson for the next meeting: Faisal Kamal
10. Next meeting to be held at N/A on N/A in N/A
9
Group Number: CEE2 Contributing Member(s): Faisal, Abdelrahman, Bryn & Sari
B1