Capital costs and energy considerations of different alternative stripper configurations for post combustion CO2 capture

Journal Identification = CHERD Article Identification = 746 Date: June 6, 2011 Time: 11:29 am

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chemical engineering research and design 8 9 ( 2 0 1 1 ) 1229–1236

Contents lists available at ScienceDirect

Chemical Engineering Research and Design

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apital costs and energy considerations of differentlternative stripper configurations for post combustion CO2

apture

ehdi Karimi, Magne Hillestad ∗, Hallvard F. Svendsenepartment of Chemical Engineering, NTNU, 7491 Trondheim, Norway

a b s t r a c t

Capturing and storing the greenhouse gas carbon dioxide produced by power plants and chemical production plants

before it is emitted to the atmosphere will play a major role in mitigation climate change. Among the different

technologies, aqueous amine absorption/stripping is a promising one. In this study, five different configurations for

aqueous absorption/stripping have been compared with regards to capital investment and energy consumption. The

process simulations are made with the use of Unisim Design and ProTreat, while for the cost calculations, data from

Turton et al. (2009) and Sinnott and Towler (2009) are used.

We cannot identify one single configuration to be the optimum always for all situations, as it depends on many

parameters like energy and material costs, interest rate, plant complexity, etc. With the assumption and estimated

parameters in this study we find that vapor recompression configuration is the best configuration because it has

the lowest total capture cost and CO2 avoided cost. In addition, the plant complexity does not increase very much

compared to the benchmark. The split-stream configuration with cooling of semi-lean amine is the second best.

However, this configuration increases the investment cost and plant complexity significantly.

The effect of heat integration between the compression section and the stripper is also considered. We can reduce

heat requirement by heat integration, but since the inlet temperature to the compressors become higher, the com-

pression efficiency will decrease and compression work will increase. In addition, the capital cost and the complexity

of the plant will increase. Because of the higher inlet temperature the water content of produced CO2 is higher and

consequently the corrosion problems is more serious in pipes and equipment for compression and injection section.

© 2011 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

Keywords: Stripper configurations; Total capture cost; CO2 avoided cost; MEA

tion of CO2 capture. Some strategies like intercooling can be

. Introduction

ossil fuels like oil, coal and natural gas are the most commonources of energy used in the world today. Large amounts ofarbon dioxide are released when fossil fuels are burned torovide power and heat or for transportation purposes. Today,bout 50% of the world’s man-made emissions of carbon diox-de come from burning fossil fuels in power plants or otherndustrial processes. Global warming is a result of increas-ng anthropogenic CO2 emissions, and the consequences wille dramatic climate changes if no action is taken. One of
he main global challenges today is therefore to reduce the
∗ Corresponding author. Tel.: +47 73 59 41 22; fax: +47 73 59 40 80.E-mail addresses: [email protected] (M. K

[email protected] (H.F. Svendsen).Received 5 November 2010; Received in revised form 18 February 2011

263-8762/$ – see front matter © 2011 The Institution of Chemical Engioi:10.1016/j.cherd.2011.03.005

CO2 emissions. Increasing the energy efficiency and a tran-sition to renewable energy will reduce CO2 emissions, butsuch measures can only lead to significant emission reduc-tions in the long-term. Carbon capture and storage (CCS) is apromising technological option for reducing CO2 emissions ona shorter time scale. Unfortunately CO2 capturing plants areenergy intensive processes. The energy consumption in theCO2 capturing plant is estimated to be 15–30% of the net powerproduction of a coal-fired power plant (Jassim and Rochelle,2006). A lot of work has been done to reduce energy consump-

arimi), [email protected] (M. Hillestad),

; Accepted 8 March 2011

effective to reduce energy requirement of CO2 capturing plant

neers. Published by Elsevier B.V. All rights reserved.

http://www.sciencedirect.com/science/journal/02638762

mailto:[email protected]



dx.doi.org/10.1016/j.cherd.2011.03.005


1230 chemical engineering research and design 8 9 ( 2 0 1 1 ) 1229–1236

Table 1 – Flue gas specification.

Temperature (◦C) 48Flow rate (kmol/h) 24,123Pressure (bar) 1.1Composition (mol fraction)

CO2 0.1186Nitrogen 0.7291Oxygen 0.0505H2O 0.1018

than the conventional configuration and it is about 4.5 and forthe stripper is about 12.

(Karimi et al., 2010). Alternative process configurations havebeen proposed to reduce operating cost and in some casescompared to capital cost of the CO2 capture process, Schachet al. (2010), Oyenekan and Rochelle (2006, 2007), Jassim andRochelle (2006), Leites et al. (2003), and Goff et al. (1996). Inaddition to steady state studies, there are some studies indynamic mode, Panahi et al. (2010) where the best controlstructure is selected base on minimizing the energy loss withconstant setpoint for control variables when different distur-bances happen.

Since large scale CO2 capture plants are very expensive tobuild for research purposes, process simulation and model-ing have an important role to play for system optimizationand in evaluation of the various process alternatives. In thepresent study, different configurations have been simulatedand the investment cost, operating cost and total capture costare estimated. All configurations are simulated for 90% CO2

capture from flue gas of a 150 MW bituminous coal powerplant. Simulations are performed in Unisim Design with theAmine Fluid package. Because most of the studies for CO2 cap-ture have been done with monoethanolamine (MEA), aqueousMEA 30 wt% is considered here as a benchmark.

2. Process description

Here five different process configurations are simulated andcompared; the conventional process with a simple stripperas benchmark, split-stream, multi-pressure stripper, vaporrecompression, and compressor integration. The specifica-tions of the flue gas are given in Table 1.

The conventional configuration is the simplest configu-ration and has a large driving force for separation. In otherconfigurations the driving forces are reduced to make the pro-cess more reversible, or the operating conditions are changedto enhance the absorption and stripping. However, at the sametime the process complexity will increase. These configura-tions are described in the sequel.

2.1. Conventional process configuration

In Fig. 1, the flow diagram depicts the conventional processstructure, including absorber, stripper, cross heat exchanger,cooler, pumps and compressors to compress CO2 up to 110 bar.

The number of segments or the packing height of theabsorber and the stripper and the lean loading are optimizedto obtain minimum energy consumption. For the other con-figurations the height of the packing is the same as for theconventional configuration, whereas lean and rich loadingsare optimized for each case. The absorber and stripper pack-ing height are 7 and 15 m, respectively. The L/G ratio is about 3in absorber and about 11 in the stripper. For the conventionalconfiguration the optimum lean loading is 0.20 and the rich
loading is 0.49.
Fig. 1 – Conventional process configuration.

The captured CO2 is compressed to 110 bar in two parts. Thefirst part is a multi-stage compressor with a pressure ratio of2 (Baldwin, 2009) in each stage. The pressure is increased upto 75 bar in this part. The gas is cooled down to 30 ◦C betweenthe stages. After the multistage compressor, CO2 is liquefiedat 30 ◦C. Because the critical temperature for carbon dioxide is31.1 ◦C, and the critical pressure is 73.8 bar, we are sure in thiscondition CO2 is in liquid phase. The liquefied CO2 is pumpedto 110 bar in the second part.

2.2. Split-stream configuration

In this configuration the rich amine is split into two streams;going to two sections of the stripper after preheating with twoseparate lean amine streams as shown in Fig. 2.

Many design parameters such as the number of segments(packing height) in the various sections of absorber and strip-per, the rich amine split ratio, and the lean loading need to beoptimized to reach the minimum energy consumption. Here, acooler can be used to cool the lean amine from the top section(semi-lean amine) of the stripper before entering to the bottomsection of the absorber. Simulation is done with a cooler thatcools down the liquid to 40 ◦C and also without intercooler. Theoptimum values for the split-stream configuration are shownin Table 2. The L/G ratio in the bottom of absorber is higher

Fig. 2 – Split-stream configuration.


chemical engineering research and design 8 9 ( 2 0 1 1 ) 1229–1236 1231

Table 2 – Optimum parameters for the split-stream configuration.

Item Absorberheight (m)

Stripperheight (m)

Rich aminesplit ratio

Semi-leanloading

Leanloading

Richloading

Top Bottom Top Bottom Top Bottom

Without intercooling 5 2 5 10 0.44 0.56 0.4553 0.1943 0.4974With intercooling 5 2 5 10 0.42 0.58 0.4468 0.1944 0.4972

2

Isbata4(

2

Ila

ofl

2

WwdsTcmCntp

more heat can be utilized from compression section but thecompression work will increase compare to the first option,

Fig. 3 – Multi-pressure stripper configuration.

.3. Multi-pressure stripper configuration

n this configuration, the stripper works at different pres-ure levels. The vapor from the bottom section is compressedefore entering the upper sections. Some parameters suchs the pressure levels, height of the packing in each sec-ion are taken from Oyenekan and Rochelle (2006) and Jassimnd Rochelle (2006) where the pressure levels are 2, 2.8 and

atm. The optimum lean loading is 0.215 for this configurationFig. 3).

.4. Vapor recompression configuration

n this process configuration, a pressure drop is created in theean amine stream and the resulting vapor is recompressednd sent to the stripper as shown in Fig. 4.

A pressure drop of about 85–90 kPa is the optimum. Theptimal lean loadings are 0.192 and 0.185 before and after theash, respectively.

.5. Compressor integration

hen CO2 is compressed to a higher pressure the temperatureill increase. Normally an intercooler takes the temperatureown between the stages. But the hot CO2 can be used as aource of heat if it’s integrated to the bottom of the column.his integration has the positive and negative effect on energyonsumption. The positive is that the external heat require-ent reduces because a portion of the heat is supplied by hot

O2. The negative side is that the gas between the stages can-ot be cool down very much because it must be hotter than

he liquid at the bottom of the column. Therefore the com-ressor inlet temperature increases and volumetric efficiency

Fig. 4 – Vapor recompression configuration.

of the compressor decreases. Consequently the compressionwork will increase. For the compressor integration configura-tion, two options are simulated; stripper with condenser andwithout condenser. If stripper has condenser, the temperatureafter the first stage is not high enough to be integrated. Theintegration then starts from the second stage. For the stripperwithout condenser, the vapor goes to the compressor directlyfrom top of the stripper as shown in Fig. 5. Here the heat can beutilized after the first compression stage. In the second option,

Fig. 5 – Compressor integration configuration.



Table 3 – Estimated purchased cost of main equipment(2009) for the base case.

Item Purchased cost (M$)

Absorber 4.81Stripper 4.00Pumps 0.42Condenser 0.27Reboiler 0.53Cross heat exchanger 2.28Cooler 0.09Make up tank 0.23Compression section 5.08

because the inlet temperature to first stage is higher and thereis more water in CO2 stream.

For both options, the liquid from the stripper is preheatedby the hot gas out of each compressor stage to reduce energyrequirement. The gas between the stages is cooled down to130 ◦C.

3. Cost calculation

In this study the cost calculations are done based on data fromTurton et al. (2009) and Sinnott and Towler (2009). In order todo an economic analysis, it is necessary to estimate the cap-ital cost of major equipment. The most accurate method isto have a price list from vendors. The next best alternativeis to use cost data based on previously purchased equipmentof the same type. Another technique, sufficiently accurate forstudy and preliminary cost estimates, is using correlations andgraphs available for various types of common equipment insome references. Here, the base condition equipment pricecalculated by Eq. (1):

log10Cop = K1 + K2 log10(A) + K3[log10(A)]2 (1)

where A is the size parameter of the equipment. K1, K2 andK3 are given Turton et al. (2009) for different equipment. Thecorrelation (1) is valid on a specific range. Out of this range weuse Eq. (2) to estimate the price.

Ca

Cb=

(Aa

Ab

)n

(2)

where A is the equipment capacity, C is the purchased cost andn is a cost exponent. The estimated cost is for base equipmentmade from carbon steel and for atmospheric pressure. To cal-culate the purchased equipment cost with different materialsand different operating pressure we can use material factor(FM) and pressure factor (FP), respectively.

Cp,Eq =∑

i

FM,iFP,iCop,i (3)

A simple technique to estimate the capital cost of achemical plant is the Lang factor method. The total cost isdetermined by multiplying the total purchased cost for all themajor equipment by a constant. The Lang factor is set equalto 4.74 in this study.

Another technique that is introduced in Turton et al. (2009)is the calculating of module cost for the main equipment. Forsome equipment like heat exchangers, vessels and pumpsbare module cost is calculated from (4) and for other equip-ment like packing and compressors use Eq. (5).

CBM = Cop(B1 + B2FMFP) (4)

CBM = CopFBM (5)

CTM = 1.18

n∑i

CBM,i (6)

n∑
CGR = CTM + 0.5
i

CoBM,i (7)

where CBM is the bare module price, B1 and B2 are the con-stant, FM is the material factor, FP is the pressure factor, FBM

is the bare module factor, CTM is the cost of making small-to-moderate expansion, and CGR is the cost of completely newfacilities in which we start the construction in a gross field.

The cost of structured packing and compressors are calcu-lated from Sinnott and Towler (2009) and other equipmentsare calculated from Turton et al. (2009). The Chemical Engi-neering Plant Cost Index (CEPCI) from Chemical Engineeringmagazine (August 2010) is used to calculate the prices as of2009. The main equipment costs for the base case are shownin Table 3.

Here, all deviations from the conventional configurationwill increase the investment costs. In split-stream configura-tion, the rich and lean amine pumps, and cross heat exchangerare the items that increase the investment cost. Semi-leanamine has a high lean loading of about 0.45 compared to leanloading that is about 0.2, and more amine flow is needed tohave 90% CO2 recovery. Therefore, a bigger rich amine pumpis needed. Also, in this configuration there are two cross heatexchangers and lean amine pump and because of the flowincrease, the capacity of this equipment is larger than for thebenchmark. In multi-pressure configuration the rich aminepump and CO2 compressor are the items that increase theinvestment cost. A more powerful rich amine pump is neededto pump the rich amine to a higher pressure (4 atm comparedto 1.8 atm in the benchmark). The two first stages of the CO2

compressor is used for increasing the pressure of the stripperand these two stages must compress more gas (vapor volumedecrease from bottom to the top) compared to the CO2 streamfrom the top of the stripper and it will increase the cost. Alsoif the maximum pressure ratio is 2, Baldwin (2009), one morestage is needed for CO2 compression compared to the bench-mark. In vapor recompression configuration the flash tank forflashing the lean amine and the CO2 compressor are the mainitems that increase the investment. Here one more stage isneeded for CO2 compression compared to the benchmark. Incompressor integration configuration a bigger compressor isneeded and it will increase the investment cost.

4. Results

4.1. Accuracy

All simulation software has built in numerical convergencetolerances and these tolerances define the simulation accu-racy. In this work the accuracy of the software was improvedfrom the default values set obtain acceptable convergenceerror. A non-equilibrium stage model developed to simu-
late the multi-component multistage mass transfer process



Table 4 – Result comparison between Unisim Design andProTreat.

Software Lean amineflow(kmol/h)

Rich loading Reboilerduty (MW)

Unisim Design 76,596 0.4867 112.8ProTreat 76,745 0.4861 114.2

ePmvncatUs2iDhtrisotbDI9aaD

tfs

The interest rate (i) and project life time (n) are 10% and 30years, respectively. We can also calculate the CO2 avoided cost

perature approach of the cross heat exchanger.

ncountered in an amine treating unit is used in the Aminesroperty Package in Unisim Design. The model incorporates aodified Murphree-type vapor efficiency to account for the

arying mass-transfer rates of individual acid gas compo-ents. The stage efficiency is a function of the kinetic rateonstants for the reactions between each acid gas and themine, the physico-chemical properties of the amine solution,he pressure, temperature and the geometry of the column.nisim Design will calculate the height of the packing for eachtage depending on the packing is selected (Unisim Design,008). For validating the results, the benchmark is simulatedn ProTreat and the result compared to the results from Unisimesign. In the work have done by Luo et al. (2009) ProTreatas been validated with experimental data and the simula-ion result is in a good agreement with experimental data. Theeactions kinetics and reaction parameters for MEA and CO2

n ProTreat are from Barth et al. (1986). ProTreat is a rate basedoftware and we could define the packing and the geometryf the towers. In this simulation we selected Mellapak M250.Yype for both absorber and stripper and the columns are sizedased on 75% of flooding as gas velocity. Both the Unisimesign and ProTreat simulations used the same specifications.

n the absorber the lean amine flow rate was adjusted to have0% CO2 recovery. In the stripper, the condenser temperaturend CO2 composition at the bottom of the stripper are defineds specifications. The results of the simulations in Unisimesign and ProTreat are shown in Table 4.

If we compare the results from the simulators we can seehat the difference between the results is less than 1%. There-ore there is a good agreement between the results of the twoimulators.

Table 5 – Summary of the simulation results with two tem

Case Leanloading

Rlo

�Tmin = 5 ◦CConventional configuration 0.2071 0Split-stream without cooler 0.4262 0.1983 0Split-stream with cooler 0.4330 0.1987 0Multi-pressure 0.2154 0Vapor recompression 0.2064 0.1987 0Compressor integration 0.2068 0Compressor integration with condenser for stripper 0.2195 0

�Tmin = 10 ◦CConventional configuration 0.1986 0Split-stream without cooler 0.4437 0.1944 0Split-stream with cooler 0.4553 0.1943 0Multi-pressure 0.2152 0Vapor recompression 0.1922 0.1860 0Compressor integration 0.1986 0Compressor integration with condenser for stripper 0.2153 0

4.2. Results summary

The conventional configuration has the lowest investmentcost and is the easiest plant to operate and control. In allother cases the investment cost and complexity increase andin most of them the energy consumption decreases. The sum-mery results are shown in Table 5.

Here, �Tmin is the temperature approach in the cross heatexchanger. Two types of energy are needed in the process, elec-trical or mechanical energy for pumps and compressors andheat for the reboiler. These two types have different exergyvalue and for comparing the total energy for the different caseswe need to unify them. In this work the heat is convertedto equivalent thermodynamic work. This means how muchelectricity can be produced with the same amount of steamfor reboiler. We assume that the temperature of steam in thereboiler (TH) is 10 ◦C higher than the reboiler temperature andthat steam condenses at 40 ◦C in the turbine (TC). The totalequivalent work for the plant is then:

Weq = Qr

(1 − TC

TH

)× � + WPumps + WCompressors (8)

where TH = TC + 10 [K] and TC = 313 K The turbine efficiency (�)to produce electricity from steam is assumed to be 75%.

The total equivalent work is calculated by using Eq. (8) andthe results are shown in the last column in Table 5 as the totalenergy requirement of the capture plant.

From Table 5 is difficult to choose the optimum configura-tion because the investment cost and operating cost are notcomparable. It is better to calculate Total Operating Cost (TOC)where the Fixed Investment Cost (FIC) is converted into a con-stant series of payment for every year of the project life. Thesum of annual capital cost and annual operating cost is thetotal annual cost of capture plant. Table 6 shows the compo-sition of operating cost where which estimated using Sinnottand Towler (2009).

The annual capital cost is calculated from Eq. (9).

FICAnnual = FIC

((1 + i)n − 1)/i(1 + i)n(9)

ichading

Totalinvestmentcost (M$)

QR (kJ/kg CO2) Weq (kJ/kg CO2)

.4896 88.90 3545 924.7

.4831 97.33 3128 854.1

.4968 97.18 2964 827.4

.4890 95.06 2400 846.4

.4906 91.37 2595 838.1

.4897 102.82 1513 964.9

.4885 93.08 3326 967.5

.4904 83.98 3611 936.4

.4794 89.38 3335 890.6

.4974 89.68 3140 859.4

.4890 89.33 2592 877.8

.4914 85.56 2722 857.8

.4904 96.38 1829 984.7

.4889 88.22 3372 975.6



Table 6 – Composition of Total Operating Cost (TOC).

Variable Operating Costs (VOC)MEA 1.5 kg/metric ton of CO2

($1500/ton MEA)Steam cost $0.064/kWhElectrical energy cost $0.011/kWhMiscellaneous operating materials 2% of VCFixed Operating Costs (FOC)Local tax 1% of FCIInsurance 1% of FCIMaintenance (M) 3% of FCIOperating labor (OL) 10% of TOCLaboratory costs 20% of OLSupervision 20% of OLPlant overheads 50% of OLOperating supplies 15% of MAdministrative cost 15% of OLDistribution and marketing 0.5% TOCR&D cost 5% TOCTotal Operating Costs (TOC) VOC + FOC

and 8.39% for �Tmin = 10). The results show this configuration

(CAC) that is an interesting parameter in economical investi-gation of CO2 capture plant with Eq. (10).

CO2 Avoided Cost = COEcapture − COEreference

Ereference − Ecapture(10)

COE is the cost of electricity ($/MWh) and E is the CO2 emission(ton/MWh).

The total cost of captured (TCC) and CO2 avoided cost (CAC)are shown in Table 7 for different configuration.

The share of investment cost to the total cost of capture isvery sensitive to interest rate. For very high interest rate theconventional configuration will be the optimum because it hasthe minimum investment cost.

5. Discussion

The different configurations compared to the conventionalconfiguration increase both investment cost and complex-ity of the plant. The energy savings and investment increasecompared to the conventional configuration are shown inTable 8. The conventional configuration with �Tmin = 5 ◦C and�Tmin = 10 ◦C are considered as the benchmarks in this table.

In Table 8 we can see that the energy saving for �Tmin = 5compared to �Tmin = 10, is not the same for the different con-figurations. The reason is that the reboiler and condenserduties are not the same for the different cases. Table 9 showsthese differences for all configurations. In this table we can seethe effect of temperature approach of cross heat exchanger ondifferent economical parameters.

In Table 9, we can see that decreasing the temperatureapproach from 10 to 5 ◦C, has a big effect on the investmentcost. It means that the temperature approach is an importantparameter and a trade off is needed to find the optimum valuefor any capture plant. Because the energy saving and increasein investment cost are different for each configuration, theoptimum value for the temperature approach is different too.The total capture cost or CO2 avoided cost can be used forthis kind of optimization. With the assumption and estimatedcost and other parameters in this study we can see that totalcapture cost changes are small positive number for all config-

◦
urations. It means for all configurations �Tmin = 10 C is better.
5.1. Conventional process configuration

The conventional configuration is simulated with the tem-perature approach in cross heat exchanger of 5 and 10 ◦Cand forms the benchmarks for other configurations. Simu-lation shows that in the conventional configuration we canhave 1.25% energy saving by decreasing the temperatureapproach in the cross heat exchanger from 10 to 5 ◦C, but theinvestment cost increases about 5.86% and the TCC increase$0.72/ton CO2. Most of this increase is related to the cross heatexchanger, because the area is doubled when the temperatureapproach decreases from 10 to 5 ◦C.

5.2. Split-stream configuration

This configuration was simulated with and without coolingof the semi-lean stream to the bottom part of the absorber.Adding a cooler to the semi-lean stream has both positiveand negative effects on the investment (cooler cost increasesinvestment, but there is a cost reduction because the sizesof the condenser and reboiler decrease). In total the invest-ment does not change very much. However, it has a positiveeffect on energy saving. By adding a cooler, the investmentdecrease 0.15% compared to the case without semi-lean coolerfor �Tmin = 5 ◦C and the energy saving is about 3.13%. Fora temperature approach of 10 ◦C, the investment increaseand energy saving are 0.34% and 3.5%, respectively. Becausethere are two cross exchangers, the investment will probablyincrease more than with the other configurations when �Tmin

changes from 10 to 5 ◦C (Table 9). A negative point for this con-figuration is that the complexity is higher than for the otherconfigurations. Absorber and stripper with two different sec-tions, two cross heat exchangers, and the need to split streamsare some factors that increase the complexity and make thecontrollability of the plant more complex.

5.3. Multi-pressure

In this configuration the stripper operates at three pressurelevels, 2, 2.8 and 4 atm. This configuration has the 3rd rankin energy requirement and total capture cost (Table 5). If theenergy consumption for different parts of this configurationis compared to the benchmark, there is a huge saving in heatenergy (reboiler duty). The saving in heat energy is 32% and28% for a temperature approach of 5 and 10 ◦C, respectively.But the energy consumption for the first two stages of thecompressor that supplies the pressure for the stripper is muchhigher than the two first stages for the conventional configu-ration. Thus, despite of good saving in heat energy, the totalenergy requirement does not decrease significantly. The totalcapture cost decrease compare to benchmarks are 0.42 and0.77 $/ton of CO2 for temperature approach of 5 and 10 ◦C,respectively.

5.4. Vapor recompression

This configuration, has first rank for �Tmin = 10 and hasthe 2nd rank for �Tmin = 5, in energy requirement, and theincrease in investment is smaller compare to the other con-figurations (Tables 5 and 8). In this configuration with a smallincrease in investment (2.77% for �Tmin = 5 and 1.88% for�Tmin = 10) we can save significant energy (9.37% for �Tmin = 5

is the optimum one because it has the lowest total capture



Table 7 – Total cost of capture and CO2 avoided cost ($/metric ton of CO2).

Different configurations Total cost of capture ($/ton of CO2) CO2 avoided cost ($/ton of CO2)

�Tmin = 5 ◦C �Tmin = 10 ◦C �Tmin = 5 ◦C �Tmin = 10 ◦C

Conventional configuration 43.56 42.84 55.43 54.69Split-stream without cooler 43.43 42.73 54.08 53.74Split-stream with cooler 42.78 42.02 52.88 52.41Multi-pressure 42.79 42.42 53.17 53.18Vapor recompression 41.90 41.27 51.97 51.47Compressor integration 56.60 53.55 80.85 75.85Compressor integration with condenser for stripper 47.47 45.80 62.03 59.99

Table 8 – Energy, investment cost, total capture cost (TCC), CO2 avoided cost (CAC) of different configuration relative to theconventional configuration (in %).

Different configurations �Tmin = 5 ◦C �Tmin = 10 ◦C

Energy Investment TCC CAC Energy Investment TCC CAC

Conventional configuration – – – – – – – –Split-stream without cooler −7.63 9.48 −0.30 −1.35 −4.89 6.43 −0.25 −0.95Split-stream with cooler −10.52 9.31 −1.81 −2.55 −8.22 6.79 −1.9 −2.28Multi-pressure −8.47 6.93 −1.77 −2.26 −6.26 6.37 −0.99 −1.51Vapor recompression −9.37 2.78 −3.82 −3.46 −8.39 1.88 −3.67 −3.22Compressor integration 4.35 15.66 29.92 25.42 5.16 14.76 25 21.16Compressor integration with condenser for stripper 4.63 4.70 8.97 6.60 4.19 5.05 6.91 5.30

Table 9 – Change of energy consumption, investment cost, total capture cost (TCC) and CO2 avoided cost (CAC) when�Tmin change from 10 to 5 ◦C.

Basecase

Split-streamwithout cooler

Split-streamwith cooler

Multi-pressure

Vaporrecompression

Compressorintegration

Compressorintegrationwithcondenser

TCC change$/ton CO2

0.72 0.69 0.75 0.37 0.63 3.04 1.67

CAC change$/ton CO2

0.74 0.34 0.47 −0.01 0.50 5.00 2.04

Energyconsumptionchange by�Tmin

(kJ/kg CO2)

−11.7 −36.5 −32.0 −31.4 −19.7 −19.8 −8.1

Investmentchange by

+4.92 +7.95 +7.50 +5.73 +5.81 +6.44 +4.86

cmc

5

Tecdcdcifipnro

�Tmin (M$)

ost. In addition, the plant complexity does not increase veryuch compared to the benchmark because one flash and one

ompressor stage is added.

.5. Compressor integration

his configuration causes an increase of investment andnergy requirement for both options, and is thus not a goodonfiguration for CO2 capture. The compression efficiencyecreases when the inlet temperature decrease and a higherompression work is needed. For the option without con-enser on the stripper a lot of water vapor goes to the CO2

ompressor, and a large compressor with a high energy shafts needed. Because gas between stages is cooled down by liquidrom the bottom of the stripper, the reduction in temperatures small and the water does not separate from the CO2 and theroduced CO2 has water content much higher than what isormal. The high water content in the CO2 increases the cor-
osion problems in compressors, piping and other equipmentf injection section.
6. Conclusion

In this study five different configurations have been inves-tigated for post combustion CO2 capture of a flue gas withabout 12% CO2 on wet basis, produced by a 150 MW bitumi-nous coal power plant. These configurations are conventionalconfiguration as a benchmark, split-stream, multi-pressurestripper, vapor recompression and compressor integration.Among these configurations, vapor recompression configu-ration is the best configuration because it has the lowesttotal capture cost and CO2 avoided cost (Table 7). In addi-tion, the plant complexity does not increase very muchcompared to the benchmark. The split-stream configurationwith cooling of semi-lean amine is the second best. How-ever, this configuration increases the investment cost andplant complexity significantly. Multi-pressure configurationand split-stream without cooling are the next configurations.Compressor integration seems to be the worst case whereboth energy requirement and investment costs and conse-
quently total capture cost and CO2 avoided cost are higherthan the conventional configuration. Decreasing the temper-



ature approach in cross heat exchanger from 10 to 5 ◦C willincrease total capture cost and CO2 avoided cost in spite ofenergy saving in all configurations.

The most important parameters in economical investiga-tion of a plant are the capital cost of the plant, the operatingcost and the interest rate. Therefore the optimum configura-tion may change with the plant location.

Acknowledgements

Financial support provided through the CCERT project(182607), by the Research Council of Norway, Shell Technol-ogy Norway AS, Metso Automation, Det Norske Veritas AS, andStatoil AS is greatly appreciated.

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Capital costs and energy considerations of different alternative stripper configurations for post combustion CO2 capture