A techno-economic analysis of post-combustion CO2 capture and compression applied to a combined cycle gas turbine: Part I. A parametric study of the key technical performance indicators

International Journal of Greenhouse Gas Control 44 (2016) 26–41

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International Journal of Greenhouse Gas Control journal homepage: www.elsevier.com/locate/ijggc

A techno-economic analysis of post-combustion CO2 capture and compression applied to a combined cycle gas turbine: Part I. A parametric study of the key technical performance indicators Ahmed Alhajaj a,c,∗ , Niall Mac Dowell b,c , Nilay Shah c a

The Institute Centre for Energy, Department of Chemical and Environmental Engineering, Masdar Institute, PO Box 54224, Abu Dhabi, United Arab Emirates Centre for Environmental Policy, Imperial College London, South Kensington, London SW7 1NA, UK c Centre for Process Systems Engineering, Department of Chemical Engineering, Imperial College London, SW7 2AZ, UK b

a r t i c l e

i n f o

Article history: Received 9 July 2015 Received in revised form 18 October 2015 Accepted 27 October 2015 Keywords: CO2 capture Absorption model MEA Gas-fired power plant SAFT-VR gPROMS

a b s t r a c t In order to mitigate significant capital expenditure and parasitic energy demands associated with post combustion capture plant, many studies focused on improving its performance and efficiency through improvement in the design, integration of utilities and selection of key operating parameters (KOPs) using various key performance indicators (KPIs). In this study, an equilibrium monoethanolamine-based CO2 capture plant and compression train model was developed, validated and then used to assess the effects of KOPs on the performance of the CO2 capture and compression process applied to a 400 MWe combined cycle gas turbine (CCGT) power plant in hot countries using selected non-monetized key economic and environmental performance indicators. These were selected so as to allow performance comparisons without resorting to economic assumptions (e.g., discount rates, costs of energy), which make such comparisons difficult. The results illustrate higher compression power and dramatic increase of cooling water requirements in coolers and washing water systems in hot countries. This work elucidates the complex compromise between minimizing capital and operating expenditure indicators, and environmental impacts. It highlights the importance of considering the whole process, as opposed to simply focusing on the energy penalty associated with solvent regeneration. © 2015 Elsevier Ltd. All rights reserved.

1. Introduction CO2 capture and storage (CCS) is gaining a broad interest as a countermeasure to global warming (Edenhofer et al., 2014). In addition to increased efficiency measures, additional deployment of nuclear power and renewable energy, CCS is expected to contribute approximately 19% of the reduction of anthropogenic greenhouse gas emissions by 2035 under the 450 ppm scenario (IEA, 2011). Although there are several CO2 capture technologies (e.g., absorption, adsorption, membrane separation and cryogenic separation) that can be applied to different CO2 sources (Boot-Handford et al., 2013; Mac Dowell et al., 2010a; Rao and Rubin, 2002; Rubin et al., 2012), post-combustion capture using amine-based solvents is a promising choice owing to its technological maturity and commercial availability (Metz et al., 2005; Rao and Rubin, 2002;

∗ Corresponding author at: The Institute Centre for Energy, Department of Chemical and Environmental Engineering, Masdar Institute, PO Box 54224, Abu Dhabi, United Arab Emirates. E-mail address: [email protected] (A. Alhajaj). http://dx.doi.org/10.1016/j.ijggc.2015.10.022 1750-5836/© 2015 Elsevier Ltd. All rights reserved.

Rubin et al., 2012; Wang et al., 2011). There are, however, important drawbacks with the incorporation of post-combustion amine-based CO2 capture in a power plant. In addition to the significant capital expenditure associated with these systems, the large parasitic energy demands of the solvent regeneration and CO2 compression systems, and amine emissions to atmosphere are also of severe concern. In order to mitigate the challenges associated with post combustion capture plant, many studies focused on improving its performance and efficiency through improvement in the design, integration of utilities and selection of key operating parameters (KOPs) using various key performance indicators (KPIs). Reducing the energy penalty associated with solvent regeneration, increasing loading capacity and minimizing degradation were the main goals for designing innovative solvents in many studies (Darde et al., 2012; Freeman et al., 2010; Mangalapally and Hasse, 2011). Reducing energy penalty imposed on a power plant by the CO2 capture plant (i.e., reboiler duty and electricity consumption) or minimizing cost of CO2 avoided or improving exergy (i.e., available work) were suggested through innovative configurations of CO2 capture system (i.e., absorber with inter coolers, stripper with

A. Alhajaj et al. / International Journal of Greenhouse Gas Control 44 (2016) 26–41

Nomenclature Notations Ac cross sectional area (m2 ) a interfacial dry mass transfer area per volume of packed bed (m2 m−3 ) interfacial wet mass transfer area per volume of a packed bed (m2 m−3 ) c number of components carbon capture and storage CCS cp specific heat capacity (kJ kg−1 K−1 ) Cr compression ratio DCC direct contact cooler degree of capture DOC DOF degree of freedom G gas molar flow rate (mol s−1 ) h molar specific enthalpy (J mol−1 ) H packed bed height of a stage (m) height equivalent to a theoretical plate (m) HETP HTU height of transfer unit (m) overall height of transfer unit (m) Hog KPIs key performance indicators key operating parameter KOPs kg mass transfer coefficient for gas phase (m s−1 ) mass transfer coefficient for liquid phase (m s−1 ) kl Kog overall mass transfer coefficient (mol m−2 s−1 ) L liquid molar flow rate (mol s−1 ) m slope of equilibrium line N number of stages number of transfer units for direct contact cooling NTIDC P pressure (Pa) Q cooling or heating duty (W) molar specific entropy (J mol−1 ) S T temperature (K) molar specific volume (m3 mol−1 ) V ˙ W power of rotary equipment (W) liquid molar composition x y gas molar composition Greek symbols  ratio of the slopes of operating and equilibrium line g (i,out) outlet gas chemical potential of component i l(i,out) mec isn   lean  rich 

(J mol−1 ) outlet liquid chemical potential of component i

(J mol−1 ) mechanical efficiency isentropic efficiency density (kg/m3 ) lean loading of solvent at the inlet of the absorber (x(CO2 ,in) /x(MEA,in) ) rich loading of solvent at the outlet of the absorber (x(CO2 ,out) /x(MEA,out) ) net solvent loading ( rich −  lean )

Subscripts and superscripts cut-off pressure of the multi-stage compression cut cw cooling water fg flue gas g gas phase components: 1 = H2 O; 2 = MEA; 3 = CO2 ; 4 = N2 i in inlet j stage l liquid phase lean solvent ls

out *

27

outlet outlet assuming isentropic path

vapor-compression and split flows to/from strippers with variable pressures) (Amrollahi et al., 2011, 2012; Karimi et al., 2011; Mac Dowell and Shah, 2014; Oyenekan, 2007). Further, minimizing energy penalty and cost of CO2 avoided, and enhancing operating capacity were examined through considering different steam supply sources (i.e., natural gas auxiliary boiler (Romeo et al., 2008), internal steam flow from power plant (Lucquiaud and Gibbins, 2011; Romeo et al., 2008), auxiliary gas turbine (Romeo et al., 2008), solar collectors with thermal storage (Mokhtar et al., 2012)) in addition to storing rich solvent at peak hours and restoring it off-peak (Chalmers and Gibbins, 2007). Other researchers focused mainly on determining the optimal value of key operating parameters (KOPs) (i.e., amine lean loading, degree of capture, stripper pressure and lean solvent temperature inlet to the absorber) of retrofitting MEA-based post combustion capture plant to a supercritical coal plant using reboiler duty, liquid circulation rate, cooling duty and cost as KPIs (Abu-Zahra et al., 2007a,b). Others have cost-optimized similar KOPs of retrofitting a similar separation technology to supercritical coal plant while using piperazine promoted calcium carbonate as the solvent (Oexmann et al., 2008). Another study implemented in GAMS used a mixed integer nonlinear programming (MINLP) approach to find the cost-optimum operating parameters and dimension of the carbon capture and compression units attached to theoretical low CO2 content exhaust gas that meet different degree of capture (DOC) (Mores et al., 2012a). It is clear that previous studies have focused on determining the optimum design and operating parameters of CO2 capture plants incorporated in fossil fuel fired power plants in cold countries in Europe such as Norway, where cooling water is readily available at a temperature in the range of 5–18◦ .1 However, the majority of the economies of the Gulf Cooperation Council (GCC) countries have extremely high CO2 emissions per capita owing to the presence of energy intensive industry (Mac Dowell et al., 2011a). Further, these countries are gas-rich, and the majority of their energy is produced from the combustion of natural gas in combined cycle gas turbines (CCGT). It is therefore important to assess the effects of attaching or retrofitting a capture plant with MEA to power plants (e.g., CCGT) in hot countries such as UAE, where the cooling water temperature reaches 33 ◦ C for 5 months of the year.2 The overall objective of this work is to develop a detailed model of an amine-based post-combustion CO2 capture plant and CO2 compression, which can be used to examine the effects of KOPs on the overall system behavior using non-monetized KPIs. An important contribution of this work is the explicit consideration of the trade-offs between capital and operating costs, and environmental impacts through inclusion of additional non-monetized indictors that describe the height of the columns, power consumption in ancillary units and amine slippage. This requires concurrent consideration of the design and operation of the entire CO2 capture and compression system. The remainder of the paper is laid out as follows: it starts by presenting an equilibrium-stage model of an MEA-based CO2 capture plant and CO2 compression train developed in gPROMS. This is followed by a validation of the capture plant model. Then, effects of

1

http://www.seatemperature.org/europe/norway/bergin.htm. http://www.seatemperature.org/middle-east/united-arab-emirates/abudhabi.htm. 2

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KOPs on the performance of the CO2 capture and compression process applied to a 400 MWe CCGT power plant were assessed using selected non-monetized KPIs. Subsequently, the regional effect on the amount of cooling water required was quantified. Finally, it presents some concluding remarks.

g

g

g

out l l l out Pout (Tout , Vout , yi,j ) = Pout (Tout , Vout , xi,j ),

∀j = 1, . . ., N, g

g

Pin,j − P = Pout,j ,

2. Methodology

∀i = 1, . . ., c, (3)

∀j = 1, . . ., N,

(4)

and 2.1. Model development

g

l Tout,j = Tout,j

The proposed model of the CO2 capture process and compression train was implemented in gPROMS3 and is illustrated in Fig. 1. The model comprises a blower that receives a flue gas from a CCGT power plant and increases its pressure to overcome the pressure drop associated with the packed bed of the absorber column and direct contact cooler (DCC). This increases the temperature of the flue gas further, which then is cooled to 40–50 ◦ C inside the DCC. This flue gas is then scrubbed using an MEA solvent in a packed column (i.e., the absorber). This rich solvent is regenerated in the stripper. A more detailed description of an amine-based CO2 capture process was presented in an earlier work (Mac Dowell et al., 2010a). The scrubber located at the top of absorber is used to remove excess water in the flue gas before venting it to the atmosphere. This is used to balance the water in the system. Following its separation from the power station exhaust gas, the CO2 is compressed in a multi-stage compression train with intercoolers to the desired pipeline pressure, typically in the range of 10–15 MPa. 2.1.1. Stage model One of the most complex aspects of the proposed model is the description of the phenomena of simultaneous heat and mass transfer with chemical reaction present in the absorber column. There are a wide range of methods available for the description of these phenomena ranging from the relatively simple (equilibrium stages) to the more complex (rate based heat and mass transfer with chemical reaction kinetics) (Kenig et al., 2001). In the interest of simplicity, flexibility and speed of convergence, an equilibriumbased model has been implemented in this work. The packed height of the column is then obtained via the HETP concept (Mores et al., 2012b; Taylor and Krishna, 1993). The SAFT-VR equation of state was used to describe the thermophysical properties and reaction equilibria in this work (Chapman et al., 1989, 1990; Galindo et al., 1998; Gil-Villegas et al., 1997). The chemical reaction between the amine and the CO2 is explicitly described in thermodynamic model (Mac Dowell et al., 2010b, 2011b; Rodriguez et al., 2012). The description of the reactions at the level of the thermodynamic model appreciably simplifies the process model (Mac Dowell et al., 2010b, 2013). The governing mass transfer equations for all these columns and the junctions are the MESH equations (i.e., mass, energy, summation and heat) for the equilibrium stage outlined below (Taylor and Krishna, 1993): Material balance in in out out Gin yi,j + L(in) xi,j = Gout yi,j + Lout xi,j ,

∀i = 1, . . ., c,

∀j = 1, . . ., N

(1)

(5)

Summation equations i=c 

out xi,j = 1;

i=1

i=c 

out yi,j = 1,

∀i = 1, . . ., c, ∀j = 1, . . ., N.

g

g

g

l l out out li,out (Tout , Vout , xi,j ) = i,out (Tout , Vout , yi,j ),

(6)

i=1

Heat balance equations g

g

g

g

g

g

in l l in out Gin hin (Tin , Vin , yi,j ) + Lin hlin (Tin , Vin , xi,j ) − Gout hout (Tout , Vout , yi,j ) l l out − Lout hlout (Tout , Vout , xi,j ) + Q = 0.

∀i = 1, . . ., c,

∀j = 1, . . ., N.

(7)

2.1.2. Heat exchanger, reboiler and condenser models 2.1.2.1. Heat exchanger model. All heat transfer equipments (i.e., coolers and heat exchangers) are modeled using conservation of mass and energy and hydraulic equilibrium represented in Eqs. (1), (3), (4), (6) and (7). 2.1.2.2. Reboiler model. The governing equations for the reboiler are the MESH equations represented below: Material balance Lin xiin = Gout yiout + Lout xiout .

∀i = 1, . . ., c

(8)

The same Equilibrium and Summation relations outlined in Eqs. (2)–(6) are used here. Energy balance g

g

g

l l Lin hlin (Tin , Vin , xiin ) − Gout hout (Tout , Vout , yiout ) l l − Lout hlout (Tout , Vout , xiout ) + Q = 0.

∀i = 1, . . ., c.

(9)

2.1.2.3. Condenser model. Similarly, the MESH equations governing the condenser are presented below: Material balance Gin yiin = Gout yiout + Lout xiout .

∀i = 1, . . ., c

(10)

Eqs. (2)–(6), which represents the Summation and Equilibrium relations are also used here. Energy balance g

g

g

g

g

g

Gin hin (Tin , Vin , yiin ) − Gout hout (Tout , Vout , yiout ) l l − Lout hlout (Tout , Vout , xiout )–Q = 0.

Equilibrium relations

∀i = 1, . . ., c.

(11)

∀i = 1, . . ., c,

∀j = 1, . . ., N,

3

∀j = 1, . . ., N.

Process Systems Enterprise: http://www.psenterprise.com/.

(2)

2.1.3. Columns design 2.1.3.1. Absorber height. In the absorber, assuming a fast reaction and a dilute solvent, the mass transfer rate of the CO2 to the solvent is expressed in Eq. (12) (Graf, 2011). This accounts for diminishing driving force along the stage through the log mean difference of the mole fraction. The overall mass transfer coefficient describes the

A. Alhajaj et al. / International Journal of Greenhouse Gas Control 44 (2016) 26–41

29

Fig. 1. Carbon capture plant with multi-stage compression process flow sheet modeled in gPROMS.

efficiency of the stage in the presence of enhanced mass transfer through chemical reaction.



in Gin yCO 2 ,j

out − Gout yCO 2 ,j

=

in yCO

2 ,j

in ln(yCO

2



out − yCO

2 ,j

out ) /yCO ,j ,j

Kog a Ac H

2

∀j = 1, . . ., N

(12)

2.1.3.2. Stripper height. For the stripper column, the number of theoretical stages plays a role in the mass transfer of the CO2 from the liquid phase to the gas phase. Then the height equivalent to theoretical plate (HETP) is calculated from Eqs. (13)–(17) (Graf, 2011; Taylor and Krishna, 1993). HTUg =

ug kg a

(13)

HTUl =

ul kl a

(14)

Hog = HTUg + HTUl =

mGin Lin

(15) (16)

 ln  

HETP = Hog

(17)

−1

2.1.3.3. DCC and scrubber heights. The height of both the DCC and scrubber are obtained from the number of transfer units required to cool the flue gas temperature to the inlet temperature of the absorber in the case of the DCC and to a temperature that keep the water balanced in the case of the scrubber column. This is calculated from Eqs. (18) and (19) while the height of transfer unit is equal to 1.2 m for packed columns (Couper, 2010). This maximum value was used for the current study. g

NTUDC =

 TLM =

g

Tin − Tout

(18)

TLM g

g

l ) − (T l (Tin − Tout out − Tin )

g ln((Tin

g

l )/(T l − Tout out − Tin ))

 (19)

2.1.3.4. Column diameters. The minimum column diameters were obtained assuming a design at 90% of maximum operating capacity, which equals to a design point of 86% of flooding velocity (Kister,

1992; Kister et al., 2007; Koch-Giltsch). The first step within this approach is to calculate the flow parameter, which relates to liquid and gas flow rates and their densities. Then, the maximum operating capacity is obtained using correlations that depend on flow parameter, surface tension and viscosity. After that, the maximum design operating velocity while operating at 90% maximum operating capacity is obtained. This will define the minimum required diameter that ensures operation in the favorable efficient region. In this study, column diameters were calculated via the empirical correlations used in earlier work (Chapel et al., 1999). The larger diameter size was used while putting into considerations that liquid redistributors will be installed properly to prevent the risk of maldistribution.

2.1.3.5. Columns physical and transport properties. Physical and transport properties correlations and sources of data are summarized in Table 1.

2.1.4. Rotary equipment (pumps, blower and compressor) In addition to the energy required for solvent regeneration, further important energy sinks are the blowers and compressors required to move the flue gas through the absorption unit and to compress the CO2 for transport. The design procedure for a compression stage for both blowers and compressors is presented in Fig. 2. The first step is to obtain the pressure output of the multi-stage compressor and the single-stage blower. In the case of the blower, this depends on the pressure drop in the absorber column, which in turn depends on its height and the type of packing used. To decrease the parasitic power consumption and cost of CO2 compression, a pump is considered to be more favorable in the last stage of pressurizing the gas in the dense phase (Aspelund and Jordal, 2007; Skovholt, 1993). Thus, the outlet pressure of the compressor should be in the high-density phase to facilitate this. The multi-stage compressor has a total of 5 stages; it consists of three low pressure stages (i.e., 1–18 bar) with an intercooler and a knock-out drum taking excess water and a two high pressure stages with an after-cooler only. A dehydration unit should be included between the low and high compression stages in order to avoid hydrate formation and corrosion in carbon steel pipelines (Aspelund and Jordal, 2007; Zhang et al., 2006); however, as detailed pipeline design was not considered in this study, this unit was not included in the model.

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A. Alhajaj et al. / International Journal of Greenhouse Gas Control 44 (2016) 26–41

Table 1 Physical and transport properties correlations and sources of data. Property

Symbol

Reference

Comment

Diffusivity of CO2 in gas Diffusivity of CO2 in water Diffusivity of CO2 in loaded amine solution Loaded amine viscosity Water viscosity Physical solubility of CO2 in water Physical solubility of N2 O in water Physical solubility of N2 O in amine solution

DCO2 ,g DCO2 ,water DCO2 ,l l water HeCO2 ,water HeN2 O,water HeN2 O,l

(Fuller et al., 1966) (Versteeg et al., 1996) (Snijder et al., 1993) (Weiland et al., 1998) (Cheng et al., 1996) (Versteeg et al., 1996) (Versteeg et al., 1996) (Wang et al., 1992)

Fuller equation

Physical solubility of CO2 in amine solution

HeCO2 ,l

(Versteeg et al., 1996)

Gas phase mass transfer coefficient Liquid phase mass transfer coefficient Gas viscosity

kg kl g

(Onda et al., 1968) (Onda et al., 1968) (Pedersen et al., 1989)

Pederson model

Enhancement factor “Pseudo first order”

En

(Danckwerts, 1970)

En =

Liquid surface tension Reaction rate constant Specific wet area

 k2 a

(Vázquez et al., 1996) (Versteeg et al., 1996) (Onda et al., 1968)

DCO2 ,l = DCO2 ,water (water /l )

HeCO ,l 2

HeN O,l 2

=



0.6

HeCO

2 water

HeN O water 2

k2 DCO ,l 2 kl

Table 2 CASTOR pilot plant input values in the simulation of carbon capture modela (Abu-Zahra, 2009). Lean solvent flow rate (m3 /h) Lean solvent MEA composition (wt%) Lean solvent temperature (◦ C) Absorber inlet flue gas flow rate (Nm3 /h) Absorber Inlet flue gas molar CO2 composition (mol%) Absorber inlet flue gas molar H2 O composition (mol%) Absorber inlet flue gas temperature (◦ C) Absorber pressure (kPa) Stripper and Reboiler pressure (kPa) Temperature rich solvent out of heat exchanger (◦ C) Degree of capture (%) IMTP 50 mm specific surface area (ap ) (m2 /m3 )

23 30.4 58.8 4915 11.86 11.00b 47.3 101.3 181 100 90 107.1c

a This comprises of closed loops of the absorber and stripper without DCC and compressor. b This was based on reported water content of coal-fired flue gas as it was not reported along the pilot plant data (Abu-Zahra et al., 2007b). c Nakov et al. (2007).

raise the liquid to the height of the absorber, stripper, DCC, and Scrubber, and also to increase the pressure outlet of rich solvent to meet the input pressure of the stripper. The pump was also used to increase the pressure of dense CO2 gas exiting the last compression stage. ˙ p= W

Fig. 2. Steps to calculate power and cooling duty of compression stages. All thermophysical properties are calculated using the SAFT-VR equation of state.

l l )/) + g(z)] ˙ liq [((Pout m − Pin

p

(20)

The p is the efficiency of the pump,  is the density of the solvent ˙ liq is the solvent flow rate (kg/s) and z is the eleva(kg/m3 ), m tion change (m). The frictional losses are neglected here and it is assumed that there is no change in the velocity. 3. Model validation

The second step is to calculate the isentropic enthalpy through an isentropic compression path to a desired output pressure for each stage. Then, we correct the output enthalpy through incorporation of the isentropic efficiency, facilitating the calculation of the exit gas temperature. After which, the mechanical efficiency is applied to find the actual work of each stage. In the last step, we calculate the cooling water required and the amount of condensate by utilizing either a condenser that represents both cooler and a knock-out drum for low pressure compression stages or utilizing only a cooler only in the high pressure compression stages. The power input of the pump (Wp ) in watts is calculated using Eq. (20). This is obtained from the effect of static head needed to

The CASTOR pilot plant data was used to validate the performance of the model presented in the previous section. The design criteria of the model enumerated in Table 2 were simulated in gPROMS in order to calculate the following quantities: 1. 2. 3. 4. 5. 6.

Reboiler duty (MJ/ton CO2 ) Solvent rich loading (mol CO2 /mol MEA) Solvent lean loading (mol CO2 /mol MEA) Absorber packed height (m) Stripper packed height (m) Absorber bulk liquid phase temperature profile.

A. Alhajaj et al. / International Journal of Greenhouse Gas Control 44 (2016) 26–41

31

Table 3 Prediction of the model compared to CASTOR pilot plant data.

Absorber height (m) CO2 inlet flow rate (kg/h) CO2 captured (kg/h) Degree of capture (%) Reboiler duty (GJ/ton CO2 ) Lean solvent loading (mol CO2 /mol MEA) Rich solvent loading (mol CO2 /mol MEA) Stripper height prediction (m) Reboiler temperature (◦ C) Maximum solvent temperature (◦ C) in the absorber column a b c

Model

Experimenta

PREb

13.3 1145 1030 90% 4.428 0.2806 0.492 9.56c 119.45 69.5

18 1145 1040 90.8% 3.897 0.28 0.46 10 118.5 75

26% – – – 13.63% 0.21% 6.96% 4.38% 0.80% 7.33%

Data source (Abu-Zahra, 2009). Percent relative error %. Using 25 equilibrium stages.

Fig. 3. Temperature profile of liquid inside the absorber column for the model against experimental data.

3.1. Model validation results The model performance results are presented in Table 3 and in Fig. 3, which presents the temperature profile in dimensionless form (i.e., dividing the actual and predicted column height by the maximum height in experiment and model respectively) to suppress height prediction errors in the analysis. Table 3 includes percentage relative error (PRE) of the model against experimental data obtained from Eq. (21). %PRE =

|Model − EXP| × 100 EXP

(21)

The performance of equilibrium stage model for the description of the absorption column was evaluated. As might be expected, the equilibrium stage model embedded with lumped hydrodynamic model underestimated the required size of absorber column (i.e., 13.3 m as opposed to 18 m). This is due to the fact that lumped hydrodynamic model used in the current study cannot capture any effects of liquid maldistribution inside the pilot plant column. Thus, a more detailed non-equilibrium model (Mac Dowell and Shah, 2015) embedded with Computational Fluid Dynamic (CFD) hydrodynamic model can capture these effects, which results in a better prediction of the height of the column. However, this was not included in this study because the focus here is in the system trends under optimum conditions (e.g., Distributors are well placed inside the column). The lower absorber height prediction is also attributed to a higher CO2 capture rate in the experiment than the intended one (i.e., 90.8% instead of 90%). There is a good agreement between model prediction and experiments in most of the variables outlined in Table 3. The predicted reboiler duty was higher than the experimental value; however,

the predicted value is in good agreement with a typical reboiler duty of 3.9–4.2 (GJ/ton CO2 ) reported in simulation models and for a well designed commercial capture plant employing MEA as solvent (i.e., Fluor Econamine FG plants) (Abu-Zahra et al., 2007a,b; Chapel et al., 1999; Mac Dowell and Shah, 2013; Mariz, 1998; Singh et al., 2003). Furthermore, the relative error of 10–15% is reasonable considering the accuracy of model and measurements in the experiments. There is a discrepancy in the rich loading between the model and the experiment while capturing similar amount of CO2 . This was a result of more liquid evaporation predicted in the model associating with the assumption of water composition in the flue gas that was not reported in the pilot plant data. This led to a lower solvent molar flow rate exiting the absorber column in the model, which balances the higher CO2 loading resulting in similar amount of CO2 being captured. The absence of reported data of liquid molar flow rate exiting the absorber column and the actual water composition of the flue gas entering the absorber column made the analysis difficult, as the water content will affect the thermodynamic state of entering flue gas. The predicted absorber temperature profile for the model is shown in Fig. 3. The predicted equilibrium stage model temperature profile was not expected to match the pilot plant temperature profile as outlined in the earlier studies (Abu-Zahra, 2009; Lawal et al., 2009; Luo et al., 2009; Taylor and Krishna, 1993). However, the equilibrium stage model was expected to have a good match for the column end temperature as shown in Fig. 3 and noted by earlier studies. A more refined description of column temperature profiles could be obtained by using a more sophisticated rate-based description of the capture process, however, this is out of the scope of this work. Here, we are focused on system trends, as opposed to the detailed behavior at the individual unit level. Thus the computational efficiency afforded by the equilibrium stage model is preferred. In summary, an equilibrium steady-state model of MEA-based CO2 capture plant and compression train in which thermophysical properties were calculated using SAFT-VR was developed in gPROMS. The SAFT-VR thermodynamic model implicitly takes into account the chemical reaction of the system, which enhances the ability of the model to predict the mass transfer driven by the chemical potential gradient. The absorber column height was calculated using the two-film theory approach and the stripper column height was calculated using an HETP approach. The validation of the model has presented the ability of equilibrium stage model to predict the overall system performance represented by the selected KPIs. A better accuracy could be obtained by adding complexity to the model such as developing a rate-based model embedded with CFD hydrodynamic model in gPROMS, however, it demands more effort to develop and converge.

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4. Parametric study and regional impact case study The objective of this study was to initially assess the effects of key operating parameters (KOPs) on the performance of the CO2 capture process and compression train model applied to a case study of an exhaust gas typical of a 400 MW combined cycle gas turbine (CCGT) power plant (Bailey and Feron, 2005) in hot countries where the availability of low temperature cooling water is severely limited. Then, the region-specific climate conditions effect on the performance of the system was studied. The first objective was achieved by performing a number of simulations of the model while manipulating one degree of freedom, which represents the studied KOP at a time in the range listed in Table 4 while keeping the remainder DOFs represented by the input variables listed in Tables 4 and 5 at their default values. The ranges listed in Table 4 were chosen using a combination of literature data and engineering judgment. The data listed in Table 5 include the base plant input data, key design and operating variables that are set at their constant values in the entire study using data from the literature. The water make up into the capture plant listed in Table 5 was reduced through utilizing the condensate from the initial stages of the low-pressure compressors and through controlling the scrubber-cooler to recover the water in the flue gas. This has the effect of reducing the water consumption of the CO2 capture process – a concern in water strained areas.

Table 4 Key operating parameters range used in the simulations of the capture plant model. Variable name

Default value

Amine lean loading (mol CO2 /mol MEA) Degree of capture Stripper and reboiler pressure (kPa) Tls inlet to absorber (◦ C) Tfg after cooling (◦ C)

0.25 0.85 180 40 50

Variable range 0.2–0.36 0.45–0.95 81–301 40–50 40.5–54

Table 5 Case study of CCGT input values in the simulation of carbon capture model. CCGT flue gas flow rate (Nm3 /h) CCGT flue gas temperature (◦ C) CCGT flue gas pressure (kPa) CCGT flue gas molar H2 O composition (mol%) CCGT flue gas molar CO2 composition (mol%) CCGT flue gas molar N2 composition (mol%) Pressure of the absorber (kPa) Pressure drops in packing (kPa/m) Water makeup (kg/ton CO2 ) Rich/lean heat exchanger T (◦ C) Scrubber & DCC cooler T (◦ C) Absorber, scrubber, and DCC cooler diameter (m) Stripper diameter (m) Norton IMTP 50 mm dry packing area (a) (m2 /m3 ) CO2 outlet pressure from compressor (kPa) CO2 product content from condenser (mol%) Compressor efficiency (isn ,mec ) (%) Pumps efficiency (%) a

18,00,000a 98 101 12 5 83b 101c 0.2d 0 10 10 14.5 8 107.1e 14000f 90.4g (80,95)h 75f

Bailey and Feron (2005). 10% of oxygen composition was assumed to be nitrogen. c Assuming constant operating pressure in the absorber while taking into account the power duties associated with higher exit blower pressure assuming linear flue gas pressure drop along the column. This will ensure that flue gas exiting the absorber column has the same operating pressure and limits the need to develop complex model for the pressure drops hydrodynamics. d Calculated using correlation of the F factor, which depends on gas flow rate and density and liquid loading range considered in the current study (Koch-Giltsch). It is resulted in a pressure drop of 0.2–0.3 kPa/m in which lower term was used here. e Nakov et al. (2007). f Rao and Rubin (2002). g Iijima (1998). h Kvamsdal et al. (2007). b

It is acknowledged that amine slippage and its degradation production to the atmosphere can be minimized by applying many technologies: the core technology is to use a water wash system installed at the top of the scrubber to cool the flue gas to a temperature that minimizes its vaporization. Other promising technologies such as acid water removal, filters and demisters in addition to utilizing UV-light minimize the slippage and reverse the degradation products (i.e., nitrosamine and nitramine) back to amine (Kolderup et al., 2011). These all can be applied once a well-defined amine emission standard is defined (e.g., 0.2 ppmV). This work will quantify the effects of KOPs on the efforts needed to minimize amine slippage without taking into account specific prevention technology. The effects of KOPs on the performance of the CO2 capture plant and compression train were studied using the following nonmonetized KPIs: 1. 2. 3. 4. 5. 6.

Reboiler duty (MJ/ton CO2 ) Cooling duty (MJ/ton CO2 ) Volume of packing (m3 /ton CO2 h−1 ) Ancillary power consumption (kWh/ton CO2 ) Amine slippage (kg/ton CO2 ) Solvent flow rate (m3 /ton CO2 ).

These variables were used to determine the effect of key operating parameters on the capital cost, operating cost, and environmental impact of the CO2 capture plant. The reboiler duty represents the main energy requirement for the CO2 capture process. The ancillary power consumption highlights the energy required to run pumps, compressors and blowers in the system. The volume of the packing per ton of CO2 per hour drives the capital cost requirement of both the absorber and stripper columns. The amine emission is indicative of the potential for deleterious environmental impact associated with the potential emission of amine degradation products such as nitrosamines. The second objective was achieved by performing a number of simulations varying the cooling water temperature between 5 and 35 ◦ C in increments of 5 ◦ C in order to evaluate the effect of region-specific climate conditions on the cooling duty. The rest of the variables listed in Tables 4 and 5 were set at their default values. The amount of cooling water required (m3 /h) is calculated using Eq. (22); the maximum output temperature of water is assumed to be 40 ◦ C owing to environmental regulation of dumping hot water in the sea. The specific heat of water is assumed to be constant at 4.18 (kJ kg−1 K−1 ) (Peters et al., 2004). ˙ cw = 3.6 m

q˙ Cp (Tcw

out

− Tcw

in )

(22)

The region specific climate conditions effects on the design of compression system, which include choosing cut off pressure and exit temperature of each compression stage were examined using power and cooling duty as indicators of the system performance. Two case studies representing compression system design in two different regions (i.e., the hot climate such as Abu Dhabi in which the CO2 gas can be cooled to 50 ◦ C and the cold climate such as Norway in which the CO2 gas is cooled to 20 ◦ C) were studied. The first step in the design was to obtain the compression cut-off pressure in these two regions using Mollier chart that ensure reaching the dense phase before using a dense pump. Then, the system power and cooling duty required to pressurize the CO2 gas to a reference pressure (i.e., 140 MPa) was obtained using simulations in gPROMS of a compression system that includes multistage compression with intercooling and a dense pump.

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33

Fig. 4. Effects of varying amine lean loading in the KPIs of the capture plant with compression.

5. Results and discussion 5.1. Parametric study A number of simulations in gPROMS of the CO2 capture process and compression train model applied to a case study of an exhaust gas typical of a 400 MWe CCGT power plant in the UAE were performed. The default values listed in Tables 4 and 5 were used as the inputs of the model while manipulating one of the key operating variables at an incremental step of 1% of the range

specified in Table 5 while holding everything else constant. The model then specifies the required solvent circulation rate and height of the columns represented by the selected KPIs. The effects of each key operating parameter on all the selected KPIs are illustrated in figures in this section. 5.1.1. Effects of amine lean loading As can be seen in Fig. 4a, the reboiler duty is very high at low amine loading. This can be explained by noting the interconnectedness between components that contributes to the overall heat duty

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Fig. 5. Effects of varying degree of capture on the KPIs of the capture plant with compression.

which consists of: (1) sensible heat required to raise the solvent to the required stripper inlet temperature, which is linked to the solvent circulation rate; (2) heat of absorption required to reverse the reaction, which depends on the amount of CO2 captured and the type of the solvent used; (3) heat of vaporization, which represents the amount of steam required to maintain the driving force, which depends mainly on the partial pressure of CO2 and water vapor in equilibrium with the liquid phase (i.e., chemical potential gradient) as highlighted by Oexmann and Kather (2010). At low amine lean loading, the driving force (i.e., the chemical potential gradient) at the bottom of the stripper column is low. Thus, more energy was required to generate more steam needed to maintain the driving force. This outweighed the benefits of lower sensible

heat required with less solvent circulated into the system as noted in Fig. 4f. Increasing the amine lean loading further by circulating more solvent into the system lowered the reboiler duty until reaching a global minimum at amine lean loading of 0.31. This is due to increased solvent chemical potential gradient at the bottom of the stripper column, which reduced the amount of steam generated to maintain driving force. Increasing the amine lean loading beyond this point increased the reboiler duty again as this operating region was dominated by the sensible heat required to heat the increased solvent rate circulating into the system. A similar trend between amine loading and reboiler duty was observed in a study of MEA based CO2 capture plant attached to a coal plant in which a minimum reboiler duty was observed at amine lean loading of

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35

Fig. 6. Effects of varying reboiler and stripper pressure on the KPIs of the capture plant with compression.

0.32 (Abu-Zahra et al., 2007b). The behavior of the cooling duty against amine lean loading followed in line with the reboiler duty (see Fig. 4b). At low amine lean loading, the condenser cooling duty, linked with more steam generated by the reboiler, outweighed the reduction in solvent cooling duty associated with less solvent circulating in the system. A similar effect of a reduction in condenser cooling duty with higher amine lean loading was observed until it reached a global minimum at an amine lean loading of 0.30. Increasing the amine lean loading beyond this point resulted in an increase of solvent cooling duty dominating the overall cooling duty in that region.

The required volume of packing representing the total height of both columns increases while increasing the amine lean loading as shown in Fig. 4c. This is due to decreased driving force of solvent entering the absorber column which increased the volume of packing needed to maintain an 85% DOC. A similar decrease in driving force was also observed in the stripper column as a result of solvent entering the stripper column at lower temperature while increasing the amine lean loading. This resulted in a gradual increase in power consumption from pumps and blowers, as indicated by Fig. 4d, because of higher solvent circulation rates and pressure drops having taller columns. Nevertheless, there is a noticeable

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Fig. 7. Effects of varying lean solvent temperatures in the KPIs of the capture plant with compression.

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37

Fig. 8. Effects of various flue gas temperatures in the KPIs of the capture plant with compression.

reduction in amine slippage while increasing the amine lean loading (see Fig. 4e); this can be explained from solvent entering the top of the absorber column at lower driving force and thus reducing CO2 transfer and heat of reaction rates at the top of absorber column, which in turn reduced the vaporization of the amine. To capture 85% of CO2 while varying the lean loading, the liquid circulation rate was varied as shown in Fig. 4f in addition to the increase in volume of packing shown in Fig. 4c. 5.1.2. Effects of degree of capture This study shows that the specific reboiler duty did not change with higher DOC as shown in Fig. 5a because solvent circulation rate per ton of CO2 (see Fig. 5f) and amine lean loading remained

unchanged. This is contrary to the results of earlier studies (AbuZahra et al., 2007b; Rao and Rubin, 2006), which were carried out under the assumption of fixed column height and thus increasing solvent flow rate was the remaining approach to increase DOC (Dugas, 2006; Lawal et al., 2009). This in turn increased the specific reboiler duty outlined in earlier studies. In this study, however, the volume of packing was not fixed and increased with higher DOC. There was a slight reduction in cooling duty while increasing DOC as a result of distributing cooling duty consumed in cooling fixed amount of flue gas over increased amount of CO2 captured (see Fig. 5b). As can be seen in Fig. 5c, there was a dramatic increase in the volume of packing required while capturing more than 85% of CO2 . This also led to higher power consumption mainly from the

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blower because of the increase in pressure drops in taller columns while the liquid circulation rate remained unchanged (see Fig. 5d). Furthermore, as can be seen in Fig. 5e, there was a rapid loss of amine when capturing CO2 at higher than 85% DOC; this was as a result of the flue gas exiting a taller absorber column at a high temperature, which in turn vaporized the amine. The solvent circulation rate per ton of CO2 remained unchanged while changing the DOC as shown in Fig. 5f due to increased volume of packing. 5.1.3. Effects of stripper and reboiler pressure The increase of stripper and reboiler pressure decreased both the reboiler duty and the cooling duty changing in line with the reboiler duty (see Fig. 6a and b). This is attributed to a consequent increase in operating temperature as shown in Fig. 6a, which increased the solvent chemical potential gradient in the stripper column resulting in less steam being generated to maintain the driving force. This outweighed the increase of energy required to vaporize the solvent at high pressure. The cooling duty decreased with higher stripper and reboiler pressure as a result of reduction in condenser cooling duty associated with less steam being generated in the reboiler in addition to minimization in compression cooling duty. Also, there was a slight decrease in the height of the stripper column (see Fig. 6c) as a result of increased driving force while operating the stripper at high pressure. Further, there was a rapid decrease in power consumption (see Fig. 6d) mainly from the compressor as a result of compressing CO2 exiting at higher pressure from the condenser. The reboiler and stripper, however, cannot be operated at higher than 120–125 ◦ C (i.e., 180–212 kPa) owing to concerns associated with solvent degradation. The development of degradation-resistant solvents appears to have a great potential to improve the overall performance of the capture plant. However, a recent study has shown that operating the solvent at high pressure (i.e., 350 kPa) lead to an increase in NGCC energy penalty associated with extracting high quality steam from intermediate pressure steam turbines, which outweighed the benefits of lower compression and specific reboiler duty (Lindqvist et al., 2014). Operating the stripper and the reboiler at vacuum pressure increased cooling duty by 34% and specific reboiler duty by 42% as a result of decreased solvent chemical potential gradient at low operating temperature. This demands more steam to be generated to maintain the driving force and this outweighed the decrease in the solvent sensible heat linked with lower specific heat at low operating pressure. This result is in agreement with the earlier study of Oexmann and Kather (2010). The main benefit of operating at low pressure is that it allows using low temperature heating water that might be available from the waste heat. However, this increased power consumption by 17% mainly from the pump and the compressor (see Fig. 6d) resulting from the increase in the height of the stripper column associated with driving force (see Fig. 6c) and the decrease in the pressure exiting from the condenser. The stripper and reboiler pressure did not affect slippage because there was no change in the inlet conditions to the absorber column. There was no change in the solvent circulation rate as shown in the Fig. 6f because the amount of CO2 captured and the amine lean loading were constant while changing the reboiler and stripper pressures. 5.1.4. Effects of lean solvent temperature inlet to absorber As can be seen in Fig. 7a and b, there was a slight decrease in reboiler duty at low Tls owing to the increase of driving force (i.e., chemical potential gradient) of CO2 transfer to the liquid phase in the absorber at low temperature, which lowered solvent circulation rate and hence sensible heat duty of the reboiler. A slight decrease in cooling duty at low Tls was observed as a result of decreased solvent vaporization rate in the absorber, which in turn reduced the

scrubber cooling duty, and this outweighed the increase in solvent cooling duty. As the solvent temperature increases, more solvent is vaporized and transferred to the gas phase, which in turn increased the scrubber cooling duty as shown in Fig. 8b. On the other hand, there was a gradual decrease in volume of packing while increasing the Tls due to diffusivity and reaction rate enhancements with higher temperature, which was captured by the overall mass transfer coefficient equation. This outweighed the decreased driving force of absorption at high temperature (see Fig. 7c). This shorter column led to a lower pressure drop through the column; thus there was a lower blower energy consumption as shown in Fig. 7d. However, higher Tls increased the MEA vaporization as shown in Fig. 7e. The results of decreasing solvent circulation rate shown in Fig. 7f while reducing the Tls can be explained by the increase in driving force of CO2 transfer to the liquid phase at low Tls through the mass balance of the whole absorber column. When the Tls was reduced, it increased the rich loading for the liquid output of the absorber. This led to a larger net solvent loading () while keeping lean amine loading and the degree of capture (Gy) at their constant values; thus decreasing the solvent circulation rate.

5.1.5. Effects of flue gas temperature after cooling There was slight decrease in reboiler duty at low Tfg after cooling shown in Fig. 8a as a result of increased solvent rich loading at the exit of the absorber, which in turn reduced solvent circulation rate and reboiler sensible heat duty. This can be explained through the mass balance of the whole absorber column. When the Tls was reduced, it increased the rich loading for the liquid output of the absorber. This led to a larger net solvent loading () while keeping the lean amine loading and the degree of capture (Gy) at their constant values; thus decreasing the solvent circulation rate. The cooling duty, however, was increased rapidly while operating at lower Tfg as a result of higher direct contact cooler duty needed to cool the flue gas in addition to increased scrubbing cooling duty resulting from more water movement to the gas phase (see Fig. 8b). This outweighed the reduction in solvent cooler duty linked with less solvent circulating around the system. As can be seen in Fig. 8c, there was a similar gradual decrease in the volume of packing while increasing the Tfg after cooling. This led to decreased blower power consumption needed to cover the pressure drops of the shorter column as shown in Fig. 8d. Also, it increased the vaporization and loss of the amine to the atmosphere as can be seen in Fig. 8e. There was a gradual decrease in solvent circulation rate while operating at lower Tfg as shown in Fig. 8f; the reason for that increase was explained earlier in Section 5.1.4.

250

m3 water /ton CO2

38

200 150 100 50 0

0

10

20 30 Cooling water temperature oC

40

Fig. 9. Effects of cooling water temperature on the amount of cooling water required.

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39

Fig. 10. Cut-off pressures obtained from isothermal compressor paths; the dots present the path for cold environment and the stripped lines show the path in hot climate. Mollier chart adapted from GPA (1998).

5.2. Regional impact The results of varying the cooling water temperature representing different regions of the world ranging from the cold climate in Norway to the extreme hot climate in Abu Dhabi are presented in Fig. 9. It shows that there was dramatic increase in the amount of cooling water required in hot countries where the cooling water temperature is higher than 25 ◦ C. This was a result of decreasing heat transfer force in hot countries represented by the reciprocal function describing this behavior. The vertical asymptote in this graph is the discharge temperature of 40 ◦ C set by the regulatory body. Fig. 10 shows an exponential increase in the amount of cooling water when the cooling water temperature gets closer to the discharge temperature (i.e., 40 ◦ C). It is worth noting that the cooling duty for the capture plant and compression train did not change with changes in the cooling water temperature as the design temperature for all the units were set at constant values. The order of contribution of cooling duty from larger to lower is as follows: condenser; solvent cooler; scrubber; DCC cooler; compressor intercooler. Another important factor that need to be highlighted is the applicability of different operating schemes and technologies in hot countries. For example, a lower amine slippage is anticipated while operating the scrubber at 40 ◦ C temperature (Mertens et al., 2012, 2013). This type of operation will not be feasible in hot countries because it will increase the amount of cooling water needed in the washing-water system at the top of the absorber column exponentially as it brings the operation closer to the vertical asymptote explained earlier. Further, the availability of chilled water might limit the potential of utilizing ammonia-based post combustion capture plants in warm and hot countries. For example, it has been shown that although the energy penalty of the ammonia based capture plant is lower than that of an MEA-based capture plant at chilled water temperature of 10 ◦ C, the ranking was the opposite at a cooling water temperature of 20 ◦ C (Linnenberg et al., 2012). The cooling water temperature was observed to have a clear impact on the compression duties as shown in Fig. 10. Assuming isothermal compression (see Fig. 10), the pressure output of the compressor needed to reach this high density depends on the

output temperature of the gas, which depends on the cooling water temperature available within the specific region of interest. Fig. 10 highlights the effect of the cooling water temperature in two different regions, one with a hot climate such as Abu Dhabi and the other one in a cold climate such as Norway. Initially, the inlet pressures of the compressor in both regions are the same but the temperature is different. As we pressurize the CO2 along the isothermal path (i.e., 20 ◦ C and 50 ◦ C), the density of the gas at the same output pressure in the hot countries is lower than that in cold countries. This is particularly important as a significant amount of CCS deployment is expected to be in the GCC countries. Thus, a higher output pressure is required in the hot countries to reach similar density obtained in cold countries at lower pressure. The discharge pressures in the dense phase are shown to be around 7 MPa (i.e., 810 kg/m3 density) in the cold environment (i.e., cooling CO2 gas to 20 ◦ C) and around 14 MPa (i.e., 670 kg/m3 density) in the hot climate (i.e., cooling CO2 gas to 50 ◦ C). In cold countries, booster pumps with lower energy consumptions and capital cost can be used to increase the pressure of the dense gas from 7 to 14 MPa. This implies an overall lower energy consumption and thus capital cost for CO2 pressurization to 14 MPa in cold countries when compared to hot countries. The simulation of this compression system resulted in a compression power of 104.5 kWh/ton of CO2 in hot countries compared to a total power of 82.4 kWh/ton of CO2 in the cold countries. The overall cooling duty of the compression in hot and cold regions were 689.9 MJ/ton CO2 and 686 MJ/ton CO2 respectively. The cooling duty required in cold country was similar to the one in hot country because of the lower compressor cut off pressure in cold countries which overcame the increase in cooling duty required to cool the flue gas to a lower temperature. 6. Conclusion A mathematical model of an MEA-based post-combustion capture plant with compression train was proposed and implemented in gPROMS. The model was then used to study the effect of key operating parameters on the performance indicators relevant to the capital cost, operating cost, and environmental impact of the CO2 capture plant applied to a case study of an exhaust gas typical

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of a 400 MWe CCGT power plant in hot countries where the availability of low temperature cooling water is severely limited using selected non-monetized KPIs. It was observed that the optimum amine lean loading that minimized the specific reboiler duty, which represents the balance between sensible heat needed to heat the solvent circulated into the system and the heat of vaporization linked to the steam generated to maintain the driving force of CO2 transfer in the stripper column, was obtained at value of 0.31. This however, required a higher volume of packing linked with more solvent circulating around the system and hence higher blower power consumption linked with higher pressure-drops in a taller column. The DOC did not affect the specific reboiler duty (unlike results obtained in earlier studies) because solvent circulation rate per ton of CO2 and amine lean loading remained unchanged while varying volume of packing. Further, there was dramatic increase in volume of packing at higher than 85% DOC leading to an increase in the blower power consumption leading to bottlenecks associated with economies of scale. An increase rate of amine vaporization was also observed at high DOC as a result of gas exiting the absorber at high temperature. Operating the stripper and reboiler at high pressure reduces reboiler duty, volume of packing and power consumption of the whole capture and compression plant but it should be run at lower than degradation temperature. It was also noticed that it is not favorable to operate the MEA-based CO2 capture plant and compression train at vacuum pressure due to increased reboiler duty, power consumption and cooling duty. Further, higher compressing power was required to pressurize the low pressure gas exiting from the stripper column. It was also observed that while reducing flue gas temperature or lean amine temperature reduces amine vaporization and reboiler duty slightly, it increases the volume of packing, blower consumption and cooling duty for reducing flue gas or solvent temperature. The results of the study have also highlighted the challenges in applying post combustion capture technologies in hot countries. It indicates a dramatic increase of cooling water consumption and compression duty in regions where the cooling water temperature is higher than 25 ◦ C most of the year. Another challenge that needs to be overcome is the higher amount of cooling water needed in the washing water system to minimize the environmental impact of amine slippage. The current results highlight the importance of performing a whole system analysis on the performance of the CO2 capture plant with compression train as opposed to focusing on solely reducing the reboiler duty. Acknowledgement A.A thanks Masdar Institute for funding this research project. References Abu-Zahra, M.R.M., (Ph.D. thesis) 2009. Carbon Dioxide Capture from Flue Gas: Development and Evaluation of Existing and Novel Process Concepts. TU Delft, Netherlands. Abu-Zahra, M.R.M., Niederer, J.P.M., Feron, P.H.M., Versteeg, G.F., 2007a. CO2 capture from power plants: Part II. A parametric study of the economical performance based on mono-ethanolamine. Int. J. Greenh. Gas Control 1, 135–142. Abu-Zahra, M.R.M., Schneiders, L.H.J., Niederer, J.P.M., Feron, P.H.M., Versteeg, G.F., 2007b. CO2 capture from power plants: Part I. A parametric study of the technical performance based on monoethanolamine. Int. J. Greenh. Gas Control 1, 37–46. Amrollahi, Z., Ertesvåg, I.S., Bolland, O., 2011. Optimized process configurations of post-combustion CO2 capture for natural-gas-fired power plant—exergy analysis. Int. J. Greenh. Gas Control 5, 1393–1405. Amrollahi, Z., Ystad, P.A.M., Ertesvåg, I.S., Bolland, O., 2012. Optimized process configurations of post-combustion CO2 capture for natural-gas-fired power plant – power plant efficiency analysis. Int. J. Greenh. Gas Control 8, 1–11.

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