Natural gas combined cycle power plant modified into an O 2 /CO 2 cycle for CO 2 capture J.-M. Amann a , M. Kanniche b , C. Bouallou a, * a Centre Énergétique et Procédés (CEP), Ecole Nationale Supérieure des Mines de Paris, 60, Boulevard Saint Michel, 75006 Paris, France b EDF, Research and Development Division, Fluid Mechanics, Energies and Environment, 6 quai Watier 78401 Chatou Cedex, France article info Article history: Received 27 April 2007 Received in revised form 16 April 2008 Accepted 20 November 2008 Available online 13 January 2009 Keywords: O 2 /CO 2 CO 2 capture Aspen plus TM Air separation unit NGCC Amine scrubbing abstract Reduction of greenhouse gas emissions is becoming essential for struggling against global warming. Pri- ority has been given to sources where carbon dioxide (CO 2 ) emissions are the largest and the most con- centrated. Power plants using fossil fuels offer a great opportunity of applying CO 2 recovery processes. The O 2 /CO 2 cycle is an interesting option since CO 2 concentration in the flue gas is highly increased. This cycle has been applied to a natural gas combined cycle (NGCC) using an advanced gas turbine (GE9H). The aim of this study is to assess by simulation the energy and environmental performances of this new type of power plant. The oxygen required is produced by an air separation unit (ASU) that can deliver oxygen with a purity ranging between 85 and 97 mol.%. A CO 2 recovery process based on a cryogenic separation of carbon dioxide from inert gases has been designed and assessed. The impact of CO 2 capture has been calculated with the Aspen plus TM software. With an O 2 purity of 90 mol.% and an 85% CO 2 recovery rate, the net electrical efficiency reaches 51.3% (based on the low heating value (LHV)). This corresponds to an efficiency loss of 8.1%-points in comparison with the base case. The quantity of avoided CO 2 is about 280 g kW 1 h 1 . These results have been compared with a conventional amine scrubbing applied to a NGCC. With a lean CO 2 loading of 0.16 mol CO 2 /mol amine, this process leads to a net electrical efficiency of 49.1% (LHV). The conversion into an O 2 /CO 2 cycle seems to be more efficient than amine scrubbing but more difficult to implement because of the specific gas turbine. Ó 2008 Elsevier Ltd. All rights reserved. 1. Introduction It is well known that carbon dioxide plays a dominating role in the greenhouse effect. The rise of CO 2 concentration in the atmo- sphere, whose origin is mainly anthropogenic, is mostly due to the combustion of hydrocarbons. Renewable energies (wind, bio- mass, solar thermal ...), alone, will not be able to provide the world with the essential energy needs, at least in the medium term. To- day, the electricity cost with renewable energies are much higher than with hydrocarbons [1], which results in the development of several research programs with respect to CO 2 capture. The activities focus on locations where CO 2 emissions are relatively high and concentrated, such as power plants, refineries, cement plants and steel industry. Power plants, except the nuclear ones, produce relatively high level of carbon dioxide. According to the US DOE (Department of Energy), around 40% of the anthropogenic CO 2 emissions in the United States are due to the combustion of fossil fuels for electric- ity production [2]. The increasing use of coal to the detriment of natural gas will contribute to raise this percentage. Improving thermal efficiencies can to a certain extent contribute to reduce CO 2 emissions. However, it is essential to develop, in the same time, processes allowing to reduce carbon dioxide emissions at low cost. Post-combustion capture thanks to an aqueous solution of amines is one of the most studied solution [3–7]. This type of sol- vent takes advantage of chemical absorption by amines (e.g. Mon- oethanolamine (MEA), N-methyldiethanolamine (MDEA),...) which allows low CO 2 partial pressure in the flue gas. However, this pro- cess is still too expensive. Recent projects for CO 2 capture in typical pulverized coal-fired power plants have reported costs as high as 60$ per avoided ton of CO 2 [8–10]. For a natural gas combined cy- cle (NGCC), some cost discrepancies exist, from 55 to almost 120 $ per avoided ton of CO 2 [8,11]. Therefore, other ways of CO 2 recovery should not be neglected. Particularly, cryogenic processes can be used for flue gas with high concentration of carbon dioxide. Such processes can be applied to oxy-fuel combustion which leads to high concentration of CO 2 in the flue gas thanks to the absence of inert gases like nitrogen. The cost per ton of CO 2 recovered would reach 25 $/ton [12]. For the oxy-fuel cycles, there are several concepts where CO 2 is used as the working fluid, including the O 2 /CO 2 cycle [6,7,13–15], the MATIANT cycle [16,17], the COOPERATE cycle [18,19] and the 0196-8904/$ - see front matter Ó 2008 Elsevier Ltd. All rights reserved. doi:10.1016/j.enconman.2008.11.012 * Corresponding author. Tel.: +33 140519111; fax: +33 146342491. E-mail address: [email protected](C. Bouallou). Energy Conversion and Management 50 (2009) 510–521 Contents lists available at ScienceDirect Energy Conversion and Management journal homepage: www.elsevier.com/locate/enconman
Transcript
Energy Conversion and Management 50 (2009) 510–521
Natural gas combined cycle power plant modified into an O2/CO2 cyclefor CO2 capture
J.-M. Amann a, M. Kanniche b, C. Bouallou a,*
a Centre Énergétique et Procédés (CEP), Ecole Nationale Supérieure des Mines de Paris, 60, Boulevard Saint Michel, 75006 Paris, Franceb EDF, Research and Development Division, Fluid Mechanics, Energies and Environment, 6 quai Watier 78401 Chatou Cedex, France
a r t i c l e i n f o
Article history:Received 27 April 2007Received in revised form 16 April 2008Accepted 20 November 2008Available online 13 January 2009
Keywords:O2/CO2
CO2 captureAspen plusTM
Air separation unitNGCCAmine scrubbing
0196-8904/$ - see front matter � 2008 Elsevier Ltd. Adoi:10.1016/j.enconman.2008.11.012
Reduction of greenhouse gas emissions is becoming essential for struggling against global warming. Pri-ority has been given to sources where carbon dioxide (CO2) emissions are the largest and the most con-centrated. Power plants using fossil fuels offer a great opportunity of applying CO2 recovery processes.The O2/CO2 cycle is an interesting option since CO2 concentration in the flue gas is highly increased. Thiscycle has been applied to a natural gas combined cycle (NGCC) using an advanced gas turbine (GE9H). Theaim of this study is to assess by simulation the energy and environmental performances of this new typeof power plant. The oxygen required is produced by an air separation unit (ASU) that can deliver oxygenwith a purity ranging between 85 and 97 mol.%. A CO2 recovery process based on a cryogenic separationof carbon dioxide from inert gases has been designed and assessed. The impact of CO2 capture has beencalculated with the Aspen plusTM software. With an O2 purity of 90 mol.% and an 85% CO2 recovery rate,the net electrical efficiency reaches 51.3% (based on the low heating value (LHV)). This corresponds to anefficiency loss of 8.1%-points in comparison with the base case. The quantity of avoided CO2 is about280 g kW�1 h�1. These results have been compared with a conventional amine scrubbing applied to aNGCC. With a lean CO2 loading of 0.16 mol CO2/mol amine, this process leads to a net electrical efficiencyof 49.1% (LHV). The conversion into an O2/CO2 cycle seems to be more efficient than amine scrubbing butmore difficult to implement because of the specific gas turbine.
� 2008 Elsevier Ltd. All rights reserved.
1. Introduction
It is well known that carbon dioxide plays a dominating role inthe greenhouse effect. The rise of CO2 concentration in the atmo-sphere, whose origin is mainly anthropogenic, is mostly due tothe combustion of hydrocarbons. Renewable energies (wind, bio-mass, solar thermal . . .), alone, will not be able to provide the worldwith the essential energy needs, at least in the medium term. To-day, the electricity cost with renewable energies are much higherthan with hydrocarbons [1], which results in the development ofseveral research programs with respect to CO2 capture. Theactivities focus on locations where CO2 emissions are relativelyhigh and concentrated, such as power plants, refineries, cementplants and steel industry.
Power plants, except the nuclear ones, produce relatively highlevel of carbon dioxide. According to the US DOE (Department ofEnergy), around 40% of the anthropogenic CO2 emissions in theUnited States are due to the combustion of fossil fuels for electric-ity production [2]. The increasing use of coal to the detriment ofnatural gas will contribute to raise this percentage. Improving
ll rights reserved.
+33 146342491.allou).
thermal efficiencies can to a certain extent contribute to reduceCO2 emissions. However, it is essential to develop, in the sametime, processes allowing to reduce carbon dioxide emissions atlow cost.
Post-combustion capture thanks to an aqueous solution ofamines is one of the most studied solution [3–7]. This type of sol-vent takes advantage of chemical absorption by amines (e.g. Mon-oethanolamine (MEA), N-methyldiethanolamine (MDEA),...) whichallows low CO2 partial pressure in the flue gas. However, this pro-cess is still too expensive. Recent projects for CO2 capture in typicalpulverized coal-fired power plants have reported costs as high as60$ per avoided ton of CO2 [8–10]. For a natural gas combined cy-cle (NGCC), some cost discrepancies exist, from 55 to almost 120 $per avoided ton of CO2 [8,11].
Therefore, other ways of CO2 recovery should not be neglected.Particularly, cryogenic processes can be used for flue gas with highconcentration of carbon dioxide. Such processes can be applied tooxy-fuel combustion which leads to high concentration of CO2 inthe flue gas thanks to the absence of inert gases like nitrogen.The cost per ton of CO2 recovered would reach 25 $/ton [12]. Forthe oxy-fuel cycles, there are several concepts where CO2 is usedas the working fluid, including the O2/CO2 cycle [6,7,13–15], theMATIANT cycle [16,17], the COOPERATE cycle [18,19] and the
J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521 511
COOLENERG cycle [20]. Zhang and Lior [21,22] proposed systemsintegrated with liquid natural gas cold exergy utilization. Thereare other concepts using H2O as the working fluid including theWater cycle [23,24] developed by clean energy systems (CES),the Graz cycle [25], and others [26]. The oxy-fuel cycles requireshuge quantities of oxygen. To reduce the penalty due to oxygenproduction, new technologies have been developed, such as chem-ical-looping combustion [27–29] and the AZEP concept [30,31].Kvamsdal et al. [6,7] made a quantitative comparison of variousoxy-fuel combustion cycles with respect to plant efficiency andCO2 emissions. Their results show that the adoption of these newtechnologies promises improved performance because they re-quire no additional energy for oxygen production, but they are stillunder development.
Captured CO2 can be used for other applications such as en-hanced oil recovery (EOR) and enhanced gas recovery (EGR) orcan be stored definitively in underground reservoirs. In term ofoil and gas productivity, ya Nsakala et al. [32] recall that CO2 istwo to three times more effective than steam for EOR and EGR.
The objective of this study is to evaluate the technical perfor-mances of a NGCC converted for oxy-fuel combustion using theO2/CO2 cycle. Simulations have been carried out using the processsoftware Aspen PlusTM. The studies made on this cycle found in theliterature give only the basic trends of this new type of powerplant. In most cases, the power plant or the recovery process arevery simplified. To take into account all the characteristics of apower plant, the flowsheet in this study is based on EDF (a Frenchelectricity company) models representing a NGCC power plantusing the GE 9H gas turbine. The oxygen required for the fuel com-bustion is supplied by an air separation unit (ASU). A particularattention has been paid to this unit because it implies a high en-ergy penalty. A high ratio of the flue gas is sent back to the combus-tion chamber of the gas turbine to control the flame temperature.Therefore, a small part of the flue gas has to be treated for CO2
recovery, which allows a reduction of the process size. The highCO2 content in the flue gas allows a capture by a cryogenic process.A sensitivity study has been made on the oxygen purity and theCO2 recovery rate to assess their impact on the net power plantoutput. This new type of power plant has been compared with apost-combustion capture by an aqueous solution of MEA.
Compressor outlet pressure MPaTurbine inlet temperature KTurbine outlet temperature K
Turbine outlet pressure MPaHP steam pressure MPaIP steam pressure MPaLP steam pressure MPaCondensation pressure PaHP steam overheating temperature KIP steam overheating temperature K
2. Description of the different concepts
The different models were developed within the Aspen PlusTM
process software based on energy and mass balances [33,34]. Therelative error tolerance for convergence was set at 0.01%.
2.1. Description of the base case
The EDF model has been built from an existing power plant: theintegrated gasification combined cycle (IGCC) of Puertollano(Spain) with a conventional gas turbine: the Siemens/KWU V94.3turbine. Simulations of the power plant have been validated withconstructor data for both natural gas and synthesis gas. This NGCCmodel has been modified to work with the advanced GE9H gasturbine, using steam for turbine blades cooling rather than air usedin conventional gas turbines. This technology allows an inlettemperature of 1700 K. The net electrical efficiency is about59.5% (based on the low heating value (LHV)). The CO2 emissionsare about 339 g kW�1 h�1.
Among the changes, the gas turbine allows a higher compres-sion ratio and a higher turbine inlet temperature. The steam cyclehas also been modified. The locations of the super-heaters of thehigh pressure steam and the intermediate pressure steam havebeen modified in the heat recovery steam generator (HRSG) to takeadvantage of the higher temperature of the flue gas. The pressureand reheat temperature levels have been increased.
The natural gas is initially at 1.953 MPa and 294.8 K. Its compo-sition is given in Table 1. It is burned in the combustion chamber(CC) with compressed air. The outlet pressure of the air compressoris set to conserve the reduced flow QR entering the gas turbine (GT)
QR ¼Q �
ffiffiffi
Tp
P
where Q (kg s�1) is the mass flow rate entering the gas turbine, P(Pa) stands for the inlet pressure of the turbine and T (K) representsthe inlet temperature of the turbine. The flue gas remains at hightemperature after being expanded in the turbine and is then usedto feed the steam cycle in the heat recovery steam generator. Thesteam is raised to three pressure levels. The high pressure steamand the intermediate pressure steam are overheated. The low pres-sure steam is heated at a 470.5 K and then mixed with the low pres-sure steam recovered at the exit of the intermediate pressureturbine. At the low pressure turbine exit, the condensed steam ispumped and heated before being recycled to the feed-water tank.The water is close to its boiling point at 0.5729 MPa. The character-istics of the combined cycle are summed up in Table 2.
2.2. Description of the O2/CO2 cycle
The O2/CO2 cycle has been built from the base case. A sketch ofthe converted cycle is shown in Fig. 1. The power plant can be
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decomposed in four main units: the air separation unit, the gas tur-bine, the heat recovery steam generator and the CO2 recovery pro-cess. The characteristics of the cycle are summed up in Table 2.
The RK-SOAVE equation of state [35] has been used for the airseparation unit, the gas turbine, the heat exchangers (for the fluegas only) in the heat recovery steam generator and the CO2 captureprocess. The STEAMNBS data system (pure water) has been usedfor the steam cycle. This thermodynamic model is based on the1984 NBS/NRC steam table correlation for thermodynamic proper-ties and on the International Association for Properties of SteamIAPS for the transport properties.
2.2.1. Air separation unitAn air separation unit has been chosen to produce the high
quantity of oxygen because, even if oxygen transport membranesare a more effective way of producing pure oxygen, they are stillunder development. It is not yet suited for large scale powergeneration. Some researches need to be made to scale them toindustrial size [8]. The air separation unit recovers oxygen fromair (ISO conditions). The oxygen flow, whose O2 purity varies from85 to 97 mol.%, is injected in the combustion chamber of the gasturbine with the natural gas and the recycled flue gas. The air flowentering the air separation unit is fixed to obtain a 5 mol.% excessof oxygen (wet base) in the outgoing flue gas of the combustionchamber. This ensures a complete combustion of the natural gas.
The air separation unit developed within Aspen PlusTM (Fig. 2) isquite conventional [36,37], even if the configuration could some-what differ from an air separation unit to another. It consists ona double distillation column which produces liquid O2 at0.15 MPa. This column has a high pressure zone which roughlyseparates oxygen from nitrogen and argon and a low pressure zonewhich finalizes the purification of the oxygen stream. Between thetwo zones of the column, a spray-condenser is used to condensatethe gaseous stream in the high pressure part to supply heat to the
Fig. 2. Air separ
low pressure part. This heat is used to evaporate nitrogen and ar-gon dissolved in the liquid oxygen stream, these species beingmore volatile than oxygen. In the flowsheet, the column has beensplit in two blocks: a high pressure column (HP column) and alow pressure column (LP column).
The air (1) supposed at ISO conditions (0.1013 MPa, 288.15 K,60% relative humidity) is compressed in a two-stage inter-cooledcompressor (CPR1). Water, hydrocarbons and acid gases (3) arethen removed by a filter. The air stream (4) is split in two streams.The stream (8) is sent to the first heat exchanger (ECH1) where it iscooled to its dew point by the N2 stream (18) recovered at the topof the low pressure column. The second air stream (5) is com-pressed and cooled before entering a second heat exchanger(ECH2) where it is entirely liquefied (7). The cooling is suppliedby the N2 stream coming from ECH1 (19) and the O2 stream (15)recovered at the bottom of the low pressure column. The two airstreams (7 and 9) are sent to the high pressure column. The expan-sion of the high pressure stream provides some cooling duties tothe system. A stream enriched with O2 (10) is recovered at the bot-tom of the high pressure column and a stream rich in N2 (11) at thetop. These two liquid streams are sent to a third heat exchanger(ECH3) where they are cooled by the N2 stream (17). They are thensent to the low pressure column to be purified. The head pressureof the low pressure column has been set to 0.15 MPa to avoid airinfiltrations. A pump is used to raise the pressure of the liquid O2
stream recovered at the bottom of the LP column. This allows com-pressing O2 at low cost. This stream is then vaporized in the heatexchanger ECH2 (16). Temperatures of the N2 stream (20) andthe O2 stream (16) have been fixed at 298.15 K at the exit ofECH2. The air flow rate (5) is fitted to reach this condition. Thetemperature and pressure levels inside the air separation unit aregiven in Table 3 for an oxygen purity of 90 mol.%.
The operation of the spray-condenser requires that the boilingtemperature of the liquid oxygen stream in the low pressure col-umn is lower than the dew temperature of the nitrogen streamat the top of the high pressure column. The pinch temperaturehas been fixed at 0.75 K. Since the pressure has been set at0.15 MPa at the top of the low pressure column, it is necessary tohave a sufficient pressure in the high pressure column to fulfill thiscondition. Fig. 3 represents the bubble curve of oxygen and nitro-gen. For an oxygen pressure of 0.15 MPa, the nitrogen pressuremust be equal to 0.53 MPa to fulfill the temperature approach ofthe spray-condenser. In practice, the streams are a blend of compo-nents. The oxygen stream boiling temperature increases with itsconcentration in oxygen. This means that the high pressure columnpressure must be increased with oxygen purity. This pressure levelis the main parameter influencing the process performance since it
ation unit.
Table 3Temperature and pressure levels inside the air separation unit for an oxygen purity of90 mol.%.
Fig. 3. Nitrogen and oxygen vapor pressure curves [38].
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acts on the outlet pressure of the air compressor and thus on itselectrical consumption.
2.2.2. Combined cycleThe compressor of the gas turbine is originally used to compress
air before the combustion chamber. For the O2/CO2 cycle, it is usedto compress the recycled flue gas. To control the flame tempera-ture, a high ratio of the flue gas must be recycled towards the com-bustion chamber via the gas turbine compressor. The combustionchamber should remain quite conventional if part of the flue gasis recycled to limit the flame temperature. The NOx formation isalso highly reduced because of the absence of nitrogen in thecombustion system [39]. For the combustor, we have used theGibbs free energy minimization, assuming the thermodynamicalequilibrium of the flue gas.
The nominal outlet pressure of this compressor and the turbineinlet temperature are kept similar to the respective base casevalues. The pressure ratio of the gas turbine compressor is howeveradjusted around nominal value to conserve the reduced flow QR
entering the expander of the gas turbine.Due to the difference of the thermo-physical properties of this
type of flue gas, the temperature level at the outlet of the compres-sor and the turbine will be different. For a given pressure ratio inthe compressor, its outlet temperature will be lower for a stream
rich in CO2 than for air. On the contrary, the turbine outlet temper-ature will be higher for a stream rich in CO2 than for air. However,the lifespan of such turbine will be almost the same as a conven-tional turbine since the revolution number of the turboshaftengines will be reduced for an operation with a gas rich in CO2
[13]. Special care should be however given to corrosion problemwhich was not studied here. The turbine outlet temperature varieswith the purity of the oxygen flow. It is as high as the O2 stream ispure. This is due to a lower dilution of the flue gas with residual N2
and Ar. The heat recovery steam generator was optimizedaccording to the turbine outlet temperature of the base case.Therefore, in this O2/CO2 cycle, the flue gas has been cooled beforeentering the boiler. The extra heat is used to warm up the recycledflow of CO2 downstream of the gas turbine compressor. This allowsreducing the fuel flow rate and then the CO2 emissions. One mustkeep in mind that such gas turbines are not yet available at thepresent time. It will be necessary to develop new design adaptedto a fluid rich in CO2.
The inlet temperature of the heat recovery steam generator isconserved and there is no major change in its configuration. Themain characteristics of the heat recovery steam generator, whichwas built from constructor data, have been conserved (pressureand temperature levels, pressure drop in the heat exchangers andheat losses).
The flue gas leaves the heat recovery steam generator atapproximately 387 K and it is then cooled to 303 K to withdrawthe bulk of water. A small part of the flue gas is sent to the CO2
recovery process. The study has shown that the ratio of flue gassent to the recovery process ranges from 8.5 to 10.2% when theO2 purity decreases from 97 to 85 mol.%. The other part is recycledback to the gas turbine compressor.
Concerning the heat recovery steam generator, the thermaltransfers should be more important due to the nature of the fluegas. This unit will require accurate calculations on heat exchangersto be optimized.
2.2.3. CO2 recovery processThe aim of the O2/CO2 cycle is to concentrate CO2 in the flue gas.
This leads to a decrease of the gas flow rate sent to the capture unit.Thus, the CO2 recovery is easier and the process size can be re-duced compared with post-combustion capture. Since the fluegas has a high concentration of CO2, a cryogenic process is well sui-ted to separate CO2 from non-condensable gases. Wilkinson et al.[37,40] have proposed two CO2 recovery processes based on thedifference between the components boiling points. Our recoveryprocess is based on the work of theses authors.
The sketch of the process can be shown in Fig. 4. The flue gas iscondensed at 303 K at the exit of the heat recovery steam genera-tor to remove the bulk of water. Then the dry flue gas (1) is com-pressed to 3.5 MPa in a three-stage inter-cooled compressor(MCPR1). This pressure level is required by the water removal pro-cess using triethylene glycol (TEG) and it is also suitable for cryo-genic separation. It is important to remove water to avoid iceformation in the process. This is also required for CO2 transporta-tion in pipeline where corrosion and hydrates formation can occur.The residual molar fraction of H2O is fixed at 20 ppm (3). This flowis then mixed with two recycled streams (14) and (21). This recycleis necessary to reach the specification on the CO2 recovery rate andpurify the final CO2 stream. The resulting stream (4) enters a firstheat exchanger (MECH1) where it is partly liquefied. From thegas–liquid separator (F1), two streams are recovered: a liquidstream (19) enriched with CO2 and a gas stream (6) which containsa non negligible quantity of CO2 and some other components likeargon, oxygen and nitrogen. This gas stream is cooled in a secondheat exchanger (MECH2) to liquefy some extra CO2. The recoveryrate depends on the outlet temperature of the stream (7) since this
Fig. 4. Cryogenic CO2 recovery process.
Table 4Simulation assumptions.
O2/CO2 cycle
Turbo-machinery mechanical efficiency % 99.89
Air separation unitCompressor isentropic efficiency % 87.0Pressure loss in the heat exchangers MPa 0.01–0.02Pressure loss at the entry of the HP column MPa 0.03Pressure loss in the HP column MPa 0.005Pressure loss in the HP column MPa 0.005
Gas turbineCompressor isentropic efficiency % 88.8Turbine isentropic efficiency % 90.1Pressure loss in the combustion chamber % 2.0Heat loss in the combustion chamber % 0.2
Amine scrubbingPressure loss in the absorber MPa 0.01Pressure loss in the stripper MPa 0.01Intermediary CO2 compression isentropic efficiency % 85.0Final CO2 compression isentropic efficiency % 87.0
0
2
4
6
8
10
12
14
16
0 0.2 0.4 0.6 0.8 1X(N2), Y(N2) /molar fraction
Pres
sure
/MPa
RK-Soave (Aspen) - 223.15 KRK-Soave (Aspen) - 248.15 KRK-Soave (Aspen) - 273.15 KWeber et al. (1984) - 223.15 KWeber et al. (1984) - 248.15 KWeber et al. (1984) - 273.15 KYorizane et al. (1985) - 273.2 K
Fig. 5. Vapor–liquid equilibria of the CO2–N2 system.
514 J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521
temperature is linked to the quantity of liquefied CO2. The lowerthe temperature, the higher the CO2 recovery. A minimal tempera-ture of 223.15 K has been fixed to prevent carbon dioxide fromsolidification. A temperature approach of 5 K between the cold in-let/outlet temperature and the hot outlet/inlet temperature hasbeen chosen. To produce the cooling duties necessary to this heatexchanger, the streams (7) and (10) are respectively adiabaticallyexpanded in the gas–liquid separator (F2) and through the valve(V1). The pressure drops are calculated to fulfill the conditionson the temperature approach. A rich CO2 liquid stream (9) and agas stream mainly composed of non-condensable gases (N2, O2
and Ar) are recovered from the gas–liquid separator (F2). The gasstream supplies some cooling duties to (MECH2) and then to(MECH1). This gas stream is still at high pressure and can be ex-panded to recover some extra power. But prior being expanded,it (16) is heated to 423.15 K by using the heat released in thethree-stage inter-cooled compressor (MCPR1). Then it is expandedin a two-stage turbine with intermediate reheat. The outlet pres-sure has been chosen slightly higher than atmospheric pressure.The liquid stream (9) recovered from the gas–liquid separator(F2) contains some amount of CO2 which must be recycled to reachthe specification on the CO2 recovery rate. Before that, it suppliessome cooling duties to the two heat exchangers (MECH2) and(MECH1). For this latter, the temperature difference between thecold inlet streams (12, 15 and 22) and the hot outlet stream (5)is equal or greater than 5 K. The stream (13) is then compressed(CPR2) to be mixed with the flue gas. After the gas–liquid separator(F1), the liquid stream (19) enriched with CO2 is expanded througha valve (V2) to evaporate part of the dissolved gases. The pressuredrop is calculated to reach a 99 mol.% purity in CO2. The gas stream(20) recovered from the gas–liquid separator (F3) is recycled backafter recompression (CPR1). The liquid CO2 stream (22) suppliessome cooling duties to the flue gas in the heat exchanger (MECH1).The outlet stream (23) is then compressed to 15 MPa in a two-stage inter-cooled compressor (MCPR2). The high pressure stream(24) is cooled to 303 K to be liquefied. Many assumptions havebeen made for the different operation units (Table 4).
The Redlich–Kwong–Soave model has been used in this process.Aspen PlusTM provides the binary interaction parameters betweenCO2 and N2 determined from Knapp and Prausnitz [41]. The equi-librium curves obtained with these parameters have been com-pared with the equilibrium data of Weber et al. [42] andYorizane et al. [43] (Fig. 5). The experimental data are well repre-sented by the software for the different temperatures. Some dis-crepancies appear with the vapor data of Yorizane et al. for the
Jou et al. (1995) - 393 KEMEA insert - 393 KMEA insert - 393 KJou et al. (1995) - 313 KEMEA insert - 313 KMEA insert - 313 K
393 K
313 K
Fig. 7. Solubility of CO2 in a 30% mass MEA aqueous solution at 313 K and 393 K.
J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521 515
highest values of the molar fraction of N2. However, the recoveryprocess works at lower pressure, the errors are thus limited. Thebinary interaction parameters given by Aspen PlusTM give satisfac-tory results and were used in the different simulations.
2.3. Description of the alkanolamine process
The alkanolamine process is placed downstream of the conven-tional NGCC. The sketch of the process is represented in Fig. 6. Theflue gas is first compressed to 0.114 MPa before the absorber to off-set the pressure drop in the column. It is then washed by an aque-ous solvent containing 30 wt.% of MEA. The amine concentrationhas been limited to avoid problems of corrosion. A solvent rich inCO2 is recovered at the bottom of the absorber. It is pumped beforethe stripper. When the pressure increases, the reboiler tempera-ture at the bottom of the stripper also increases. The solvent pres-sure is thus limited by the temperature level of the low pressuresteam required for the amine regeneration. The low pressure steamis extracted from the steam cycle before the low pressure turbineat 0.32 MPa. Considering 0.1 MPa pressure loss between theextraction point and the recovery process, the steam pressure nearthe stripper is 0.22 MPa. However, only the condensation heat ofthe low pressure steam can be used in the reboiler of the stripper.The steam is cooled until its dew point. This steam is then con-densed to supply the heat required in the reboiler of the stripper.The outlet pressure of the solvent pump is chosen to have a pinchtemperature of 5 K between the steam and the reboiler. Beforeentering the stripper, the rich solvent is preheated in the heat ex-changer (ECH2) by the lean solvent recovered at the bottom of thestripper. A pinch temperature of 10 K has been selected for thisheat exchanger. The lean solvent recovered at the bottom of thestripper has the same residual quantity of CO2 as the solvent enter-ing the absorber. The CO2 gaseous stream recovered at the top ofthe stripper is compressed to 6.5 MPa in an inter-cooled compres-sor. It is then dehydrated by a triethylene glycol process. The resid-ual quantity of water has been set to 20 ppm molar. The CO2
stream is then compressed to 15 MPa and cooled to 313 K to be liq-uefied. The CO2 purity in this stream is higher than 99.9 mol.%.
Several lean CO2 loadings in the solvent have been assessed. Butonly two have been retained in this study. The first one is equal to0.25 mol CO2/mol MEA. This value lowers the energy consumptionduring the amine regeneration. This value is also the one found byAlie et al. [44]. The second one is more realistic and is equal to0.16 mol CO2/mol MEA.
The calculations are based on equilibrium data. The Aspen PlusTM
software provides some inserts for electrolyte systems. Since theresults are highly dependent on the equilibrium curves, a particu-lar attention has been paid to select a rigorous insert. The EMEA
Fig. 6. Alkanolamine CO
and MEA insert have been assessed for different temperaturesand different amine concentrations. The equilibrium curves ofthe CO2–H2O–MEA system have been compared with the data ofJou et al. [45] and those of Austgen and Rochelle [46]. For instance,Fig. 7 displays the CO2 solubility in an aqueous solution containing30 mass.% MEA at 313 and 393 K. There is good agreement for thehighest temperature whatever the CO2 loading is. For 313 K, thereare some discrepancies for the highest values of CO2 loadings.Between the two inserts, the EMEA one seems to be the best. Thisinsert matches very well the data of Austgen et al., correspondingto an amine concentration of approximately 15.5 mass.%. TheEMEA insert uses the electrolyte non-random two liquid (NRTL)model, which is suited for electrolytic aqueous solutions. It hasbeen used for the different simulations. Some assumptions havebeen made concerning the pressure loss in the absorber and thestripper (Table 3).
3. Results and discussion
Concerning the O2/CO2 cycle, a sensitivity study was carried outon the purity of the oxygen stream. Four oxygen concentrationswere selected: 85, 90, 95 and 97 mol.%. The latter concentrationwas selected in agreement with the work of Wilkinson et al. [37]who reported that, over this purity, it is necessary to separate ar-gon from oxygen. This increases the cost of oxygen production aswell as the capital expenditure of the air separation unit. The im-pact of the CO2 recovery rate was also studied. The recovery rateranges from 75 to 93.4%. The results of the O2/CO2 cycle are com-pared with those obtained for the amine process.
2 removal process.
Table 7Flue gas composition at the entry of the cryogenic CO2 recovery process.
516 J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521
3.1. The O2/CO2 cycle
3.1.1. Penalty for the air separation unitThe air separation unit has a big impact on the net efficiency of
the power plant. The air compressor is the main energy consumingcomponent in this unit. The electrical consumption is thus highlydependent on the air flow rate entering the compressor and onthe specified outlet pressure.
The performances of the air separation unit have been analyzedfor the different oxygen purities. The pressure in the high pressurecolumn must be increased with oxygen purity. The high pressurecolumn pressure must be equal to 0.421 MPa for an O2 purity of85 mol.%. But it rises until 0.538 MPa for a purity of 97 mol.%.The outlet pressure of the air compressor must be increased from0.500 to 0.617 MPa to compensate for this phenomenon. It was as-sumed that the pressure drops along the process do not vary withthe outlet pressure of the air compressor. The specific work of theair separation unit ranges from 221.3 to 268.7 kW h t�1 of the oxy-gen stream (Table 5). These values are in the range of usual electri-cal consumptions of this kind of air separation unit [5,13,47]. InTable 6, it can be observed that the specific consumption of theair separation unit is around 245 kW h t�1 O2 for an O2 purity of95 mol.%. This is slightly lower than our results.
The penalty of the air separation unit on the net efficiency hasbeen determined from the converted power plants without CO2
capture. The penalty corresponds to the difference between thenet electrical efficiency of the base case and of the O2/CO2 cycle.This penalty ranges between 5.4 and 6.3%-points depending onthe oxygen purity. If we only consider the air separation unit con-sumption, use of high oxygen purity is not favorable to the cycle.However, this purity acts on the performances of the CO2 recoveryprocess and on the maximum CO2 recovery rate.
3.1.2. Flue gas compositionThe flue gas composition is function of the oxygen purity. In-
deed, at the heat recovery steam generator exit, CO2 is all the morediluted since the O2 purity is low. The flue gas composition aftercooling at 303 K and water removal is given in Table 7. The nitro-gen concentration decreases with O2 purity whereas the concen-tration of argon is quite stable. In fact, with an O2 purity lowerthan 97 mol.%, argon is not separated from oxygen. Concerning teCO2 concentration in the flue gas, it ranges from 66.8 to85.5 mol.%. Therefore, it will be easier to capture CO2 when oxygenpurity is high.
Table 5Air Separation Unit performance with O2 purity.
O2 purity mol.% 95 90O2 flow t h�1 285 635.8ASU power MW 70 146Specific consumption kWh t�1 O2 245.6 229.6
a Purity in mass.%.
3.1.3. CO2. recovery rateThe simulations have shown that approximately 90% of the flue
gas leaving the heat recovery steam generator must be redirectedtowards the gas turbine to control the turbine inlet temperature.In fact the combustion temperature is controlled in the model tobe kept at classical level by adjusting the flue gas recirculation rate.Only a small part of flue gas have to be treated to recover CO2.
The recovery process has been optimized to lower its electricalconsumption. In the following, we will use TCOLD1 as the outlettemperature of the hot stream (5) leaving the first cryogenic heatexchanger (MECH1) and TCOLD2 as the outlet temperature of thehot stream (7) leaving the second cryogenic heat exchanger(MECH2). TCOLD2 acts on the CO2 recovery rate. The influence ofTCOLD1 on the process performances has been assessed for differ-ent values of TCOLD2. Fig. 8 represents the specific electrical con-sumption of the process for an oxygen purity of 90 mol.%. Theevolution of the specific consumption of the recovery process isnot linear with TCOLD1. For different values of TCOLD2, the mini-mum electrical consumption can be found at nearly the same valueof TCOLD1. This value has been determined for each oxygen purity.It ranges from 258.15 to 265.15 K when oxygen purity varies from85 to 97 mol.%. The CO2 recovery process has been optimized withthese values.
The performances of this process depend on the flue gas compo-sition (Table 8). The flue gas compression at the inlet of the processdepends only on the flue gas flow rate, and therefore on the O2 pur-ity. This compressor is the more consuming element in the process,
Andersson and Maksinen [47] Liljedahl et al. [5]
95 97 99a
635.8 635.8 371.8155 159 95.8243.8 250.1 257.6
140
145
150
155
240 245 250 255 260 265 270TCOLD1 (K)
Spec
ific
elec
trica
l con
s(k
Wh.
t-1 C
O2)
Fig. 8. Specific electrical consumption of the cryogenic CO2 recovery process.
Table 8Performances of the cryogenic CO2 recovery process.
J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521 517
far away from the compression of the recycled streams and the fi-nal compression of the CO2 stream. The wastes expansion powerdepends on the wastes flow rate and thus on the oxygen purityand on the CO2 recovery rate. The recycle compression power de-creases with oxygen purity since the CO2 concentration in the fluegas increases with oxygen purity. Otherwise, it increases with therecovery rate since the recycled flow rate increases. The final com-pression of the CO2 stream depends on the stream flow rate and onthe inlet pressure. Thus, for a CO2 recovery rate of 75%, the com-pression power decreases from 4.0 to 3.4 MW when oxygen purityranges from 85 to 97 mol.%. The compressor inlet pressure is as lowas the oxygen purity is small. In fact, the CO2 stream purificationrequires a higher pressure drop through the valve (V2) to reachthe specified purity when oxygen purity is low.
The maximum achievable recovery rate depends on the O2 pur-ity. For a fixed recovery rate, the more the flue gas is diluted withnon-condensable gases, the lower the minimum temperature inthe process is. Since we have specified a minimum temperatureto avoid carbon dioxide from solidification, the recovery rate islimited. It can be observed in Table 9 that, for an O2 purity of85 mol.%, the maximum recovery rate reaches 80.9% and increasesto 93.4% for a purity of 97 mol.%.
3.1.4. Net output lossFor each oxygen purity, the net electrical efficiency of the O2/
CO2 cycle has been calculated (Fig. 9). Without capture, the effi-ciency ranges from 53.3 to 54.0% (LHV), depending on the O2 pur-ity. The difference between the cases comes from the electricalconsumption of the air separation unit.
Table 9Maximum achievable CO2 recovery rate with the cryogenic process.
Fig. 9. Net electrical efficiency of the O2/CO2 cycle.
The net electrical efficiency is better for a low O2 purity. For anoxygen purity of 85 and 97 mol.%, the difference in the efficiency ofthe cycle is about 0.4%-point, although the air separation unit pen-alty is about 0.9%-point higher for the case with the purest oxygen.The recovery process is thus more efficient with higher oxygenpurity. But it is not sufficient to compensate for the penalty dueto the air separation unit. However, the maximum achievablerecovery rate is higher. The level of oxygen purity must be chosenin accordance with the targeted recovery rate and the power plantefficiency. For an 85% recovery rate and an oxygen purity of90 mol.%, the net electrical efficiency reaches 51.3% (LHV).
Fig. 10 shows the global efficiency loss in comparison with thebase case without capture. The loss does not vary much with therecovery rate since the air separation unit and the flue gas com-pression at the entry of the recovery process are independent ofthe recovery rate. The air separation unit represents about thetwo third of this loss, between 67% and 76%. The whole loss rangesbetween 7.9% and 8.7%-points according to the oxygen purity andthe recovery rate. For the two highest oxygen purities, the effi-ciency loss is quite similar. But for a purity of 90 mol.%, the effi-ciency loss is about 0.25%-point lower. The benefit of a loweroxygen purity decreases below this purity. Indeed for a purity of85 mol.%, the efficiency increases only by 0.1%-point. The penaltyof the recovery process and of the air separation unit evolve inver-sely with the oxygen purity. It seems that there is an optimal oxy-gen purity for which the global efficiency loss is minimum.However, this purity, which is lower than 85 mol.%, leads to a verylow recovery rate.
For an 85% recovery rate and an O2 purity of 90 mol.%, the cycleefficiency falls about 8.2%-points compared with the base case. Theair separation unit contributes to an efficiency loss of 5.8%-pointsand the recovery process to 2.4%-points. A higher CO2 recovery ratecan be reached by choosing a higher oxygen purity withoutdegrading significantly the net electrical efficiency of the powerplant.
3.1.5. Comparison with previous worksTable 10 compares the results of this work with those of other
authors [6,7,13–15]. Those authors have a very high CO2 recoveryrate since they did not consider a purification of the flue gas.
Fig. 10. Net electrical efficiency loss of the O2/CO2 cycle comparing with the basecase.
Table 10Performances of the O2/CO2 cycle.
This work Dillon et al. [13] Bolland and Mathieu [14]a Bolland and Undrum [15] Kvamsdal et al. [6,7]b
Type of turbine 9H 9FA 9FA GE9351FANet electrical efficiency of the base case % 59.5 56.0 55.2 58.0 56.7
CO2 recovery process penalty %-point 2.4 5.0 2.4 2.0 2.5Recovery rate % 85.0 97.2 100 100 100CO2 emissions g kW�1 h�1 59 11CO2 captured g kW�1 h�1 336 403Avoided CO2 g kW�1 h�1 279 360
a Results given for a flow of 100 kg s�1 in the gas turbine compressor.b The penalties are determined with the net efficiency of the base case whereas the authors give the penalties regarding with the gross efficiency of the O2/CO2 cycle.c The number does not take into account the O2 compression from 0.238 MPa until 3.5 MPa.d The first number correspond to the ASU delivering gaseous O2 at 0.238 MPa and the second number correspond to the O2 compression.
518 J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521
When comparing the results of the air separation unit, it can beconcluded that the specific consumption of our air separation unitis lower. However, in our study, the O2 purity has been chosen at90 mol.% to achieve a 85% recovery rate. Dillon et al. [13] havepractically the same specific consumption as us. But, the otherauthors reported higher consumption because their air separationunits produce gaseous oxygen at low pressure which is com-pressed at the outlet of the air separation unit. We have chosento produce liquid oxygen because compression is less energydemanding when performed at liquid state by a pump than atgas state by mean of a compressor. The air separation unit leadsto an efficiency loss of about 5.8%-points in our simulation, 6.3%-points for Dillon et al., 7.3%-points for Kvamsdal et al. and over7.9%-points for the other authors.
Concerning the CO2 recovery process, the penalty induced byour process (2.4%-points) is of the same order of magnitude asthose of Bolland and Mathieu [14], Bolland and Undrum [15] andKvamsdal et al. [6,7]. Their recovery process penalty is slightly low-er since there is no real flue gas purification. The flue gas is justcompressed to fulfill the transportation conditions. The energyconsumption of the recovery process of Dillon et al. [13] is signif-icantly higher. It may be partly due to the compression of the fluegas after the heat recovery steam generator. We have chosen an in-ter-cooled compression to lower the power consumption but theseauthors may have chosen an adiabatic compression resulting inhigher energy consumption.
Our net efficiency loss is lower compared with other authors.This is true even with higher oxygen purity and higher CO2 recov-ery rate. But their cycles can be optimized. Concerning Dillon et al.[13], the recovery process power is far too big. Concerning theother authors, their air separation units are not efficient since gas-eous oxygen is produced at low pressure. The O2 compression,which increases the efficiency loss, can be avoided by using anair separation unit with a pumped liquid oxygen system [36].
3.2. The base case with amine scrubbing
3.2.1. Specific consumptions in the CO2 recovery processFor the amine recovery process, the global power consumption
is due to three elements: the flue gas compressor before the absor-
ber, the steam extraction and the final compression of the CO2
stream. The power consumption of the solvent pumps is not con-sidered here since it is largely lower than the other power con-sumptions. A specific consumption has been determined for eachof these elements. Knowing the flue gas flow characteristics enter-ing the process and the recovery rate, the total power consumptioncan be calculated.
The flue gas recovered at the heat recovery steam generator exitcontains 4.97 mol.% of CO2. It is compressed to 0.114 MPa and re-quires 5.0 kWh per ton of flue gas. The heat duty in the stripper de-pends on the CO2 concentration in the flue gas and on the lean CO2
loading in the solvent. For a lean CO2 loading in the solvent of0.25 mol CO2/mol MEA, the thermal energy requirement reaches3.56 MJ kg�1 of CO2 recovered. But for a value of 0.16 mol CO2/mol MEA, which is more realistic, the thermal energy requirementincreases to 5.44 MJ kg�1. Geuzebroek et al. [48] report a value of4.3 MJ kg�1 CO2, based on field data with a lean CO2 loading of0.16 mol CO2/mol MEA and 4.9 MJ kg�1 in their simulation with alean CO2 loading of 0.21 mol CO2/mol MEA. The field data reportlower energy consumption than us. But Geuzebroek et al. reporta higher value than the field data too although their lean CO2 load-ing is higher. Our first value may be too optimistic since it assumesa high lean CO2 loading. Our second value may overestimate theheat duty. The difference with the values reported by Geuzebroeket al. may come from different operating conditions. Moreover thethermal energy requirement at low lean CO2 loading is highlydependent on the lean CO2 loading. Finally, the specific consump-tion of the final CO2 concentration has been evaluated at86.1 kW h t�1 of CO2 recovered for the highest lean CO2 loadingand at 89.7 kW h t�1 for a value of 0.16 mol CO2/mol MEA. Thethree specific consumptions do not vary much with the recoveryrate.
3.2.2. Net output lossThe results are displayed in Table 11 for the two lean CO2
loading values. The efficiency loss ranges between 7.7% and10.3%-points according to the lean CO2 loading value. The steamextraction is the main cause of efficiency reduction, even with anoptimistic value of the heat duty. It can be noticed that, for thelowest lean CO2 loading value, more than 85% of the low pressure
Table 11Performances of the NGCC power plant with post-combustion capture.
This work Davison [3] Kvamsdal et al.[6,7]
Bolland and Undrum[15]
Type of turbine 9H GE9FA GE9351FA 9FANet efficiency of the base case % (LHV) 59.5 55.6 56.7 58.0
Amine MEA(30 wt.%)
MEA(30 wt.%)
Fluor’s Econamine FG + (MEA30 wt.%)
MEA (30 wt.%) MEA
Lean CO2 loading mol CO2/molamine
0.16 0.25
CO2 recovery rate % 85 85 85–90 90 90Specific power consumption for flue gas
compressionkW h t�1 5.0 5.0 0.34a
Specific power consumption for CO2
compressionkWh t�1 CO2 89.7 86.1 91.7b
Specific heat duty in the stripper MJ kg�1 CO2 5.44 3.56 3.6 3.8Auxiliaries MWe 2.6 2.6Flue gas compression MWe 10.9 10.9Low pressure steam extracted MWe 59.3 38.8CO2 compression MWe 12.3 11.8Low pressure steam extracted % 88.7 58.0Net electrical power output % 393.2 414.2Net electrical efficiency % (LHV) 49.1 51.8 47.4 47.9 49.6Net efficiency loss %-point 10.3 7.7 8.2 8.8 8.4Efficiency loss due to flue gas compression %-point 1.4 1.4 1.7 2.0 1.8Efficiency loss due to steam extraction %-point 7.4 4.9 4.3 4.3 4.5Efficiency loss due to final CO2 compression %-point 1.5 1.5 2.2 2.3 2.0CO2 emissions g kW�1 h�1 61 58 66Avoided CO2 g kW�1 h�1 277 280 313
a Unit: MJ kg�1 CO2.b Compression until 10 MPa.
Table 12Net electrical efficiency (LHV) for amine scrubbing and for the O2/CO2 cycle.
Lean CO2 loading (mol CO2/mol MEA) Amine scrubbing O2/CO2
0.16 0.25
CO2 recovery rate (%)75 50.2 52.5 51.6a
85 49.1 51.8 51.3b
95 48.1 51.0 50.7c
a For this recovery rate, an O2 purity of 85 mol.% is required.b For this recovery rate, an O2 purity of 90 mol.% is required.c This efficiency is given for a recovery rate of 93.4% and an O2 purity of 97 mol.%.
J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521 519
steam is extracted, which will modify the isentropic efficiency ofthe low pressure turbine.
3.2.3. Comparison with previous worksDavison [3], Kvamsdal et al. [6,7] and Bolland and Undrum [15]
reported approximately the same efficiency loss, respectively 8.2%,8.8% and 8.4%-points. This is higher than our best case but lowerthan our worst case. But our CO2 recovery rate is lower. If therecovery rate is increased to 90%, our efficiency loss increases by0.4%-point for the highest lean CO2 loading and by 0.6%-point forthe lowest one. However, the efficiency loss due to the low pres-sure steam extraction is higher in our case. In our simulations,the extracted steam is condensed and returned to the feed-watertank. But before entering this tank, it is preheated to reach thespecified tank temperature. Part of the low pressure steam is ex-tracted from the steam cycle at 470.5 K to supply the heat requiredfor this operation. The ratio between the power reduction due tosteam extraction and the heat duty required in the reboiler reaches0.303 in our simulations. This is higher than the value reported byBolland and Undrum, who reported a value between 0.21 and 0.22for a low pressure steam extracted at the same pressure. The biggerefficiency loss found in our simulation may come from the extrasteam consumption for preheating the condensed steam beforethe feed-water tank.
3.3. Outcome
Table 12 compares the amine scrubbing and the O2/CO2 cycle. Itappears that converting the NGCC into an O2/CO2 cycle give prac-tically the same results than amine scrubbing for the highest leanCO2 loading. But, in practice, the lean CO2 loading seems not to belower and a more realistic value results in a higher efficiency lossthan with the O2/CO2 cycle. On the other hand, there is no availablegas turbine working with flue gas highly rich in CO2. Moreover, thethermal energy requirement in the alkanolamine acid–gas removalprocess can be lowered by modifying the configuration of the pro-cess and using advanced solvents.
It can be noticed that the O2/CO2 cycle can not reach a recoveryrate of 95% even with an oxygen purity of 97 mol.%. The CO2 recov-ery process must be modified to recover more CO2. With aminescrubbing, the maximum achievable CO2 recovery rate dependson the type of solvent. Indeed, the recovery rate is directly linkedto the thermodynamic of the system.
4. Conclusion
With the aim of evaluating the performance of a NGCC con-verted into an O2/CO2 cycle, the efficiency of the conversion wasevaluated using the Aspen plusTM process software. The power plantflowsheet was built from an existing power plant. The O2/CO2 cycleefficiency was compared with a conventional post-combustioncapture based on chemical absorption. The chemical solvent is anaqueous solution with 30 wt.% of MEA and a lean CO2 loading valueof 0.16 and 0.25 mol CO2/mol MEA.
For the O2/CO2 cycle, an air separation unit producing liquidoxygen was chosen to minimize the power consumption. This al-lows the reduction of the compression power of the oxygenstream. The specific consumption of the air separation unit rangesfrom 221.3 to 268.7 kW h t�1 O2 depending on oxygen purity. Asensitivity study on the O2 purity showed that a purity of85 mol.% was the best trade-off between the air separation unitpenalty and the recovery process penalty but the CO2 recovery rate
520 J.-M. Amann et al. / Energy Conversion and Management 50 (2009) 510–521
is limited to 80.9%. Lower oxygen purity could even be better.However, this will limit the maximum achievable CO2 recoveryrate. For a highest specification on the CO2 recovery rate, higheroxygen purity is required. For an 85% CO2 recovery rate and an oxy-gen purity of 90 mol.%, the net electrical efficiency reaches 51.3%(LHV) for the O2/CO2 cycle. This corresponds to an efficiency lossof 8.2%-points relatively to the base case. The quantity of CO2
avoided is about 280 g kW�1 h�1.The amine recovery process has three main energy consuming
elements: the flue gas compressor, the steam extraction for amineregeneration and the final CO2 compressor. The flue gas compres-sor requires around 5.0 kWh per ton of flue gas, the stripper needs3.56 MJ kg�1 CO2 for a lean CO2 loading of 0.25 mol CO2 / mol MEAand 5.44 MJ kg�1 for a value of 0.16 mol CO2 / mol MEA. The CO2
compression requires 86.1 kW h per ton of CO2 recovered for thehighest lean CO2 loading and 89.7 kW h t�1 for the lowest one.For an 85% recovery rate, the net electrical efficiency of powerplant decreases respectively by 7.7 and 10.4%-points for the twoloading cases. Steam extraction is the main energy consuming ele-ment of the recovery process.
Even if the O2/CO2 cycle displays good energetic results, theneed to design a new gas turbine is a real obstacle. Amine scrub-bing is more promising since its technology is well-known.Developments on the configuration of the amine process and thetype of solvent are still required to lower the thermal energyrequirement required for the solvent regeneration.
Acknowledgement
The authors wish to thank ARMINES and EDF (Electricité deFrance) for their financial support to this investigation.
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