Experimental Investigation on CO2 Post−Combustion Capture by

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Experimental Investigation on CO2 Post-Combustion Capture by Indirect Thermal Swing Adsorption Using 13X and 5A Zeolites Je´ roˆ me Merel, Marc Clausse,* and Francis Meunier Laboratoire de Ge´ nie des Proce´ de´ s pour l’Energie, l’EnVironnement et la Sante´ , (LGP2ES - EA 21), Cnam, IFFI, case 331, 292 rue Saint-Martin, 75141 Paris Cedex 03, France

Experimental results of CO2 post-combustion capture for a TSA process including an internal heat-exchanger (indirect heating/cooling) are presented. The comparative experimental study is carried out on 13X and 5A zeolites, with a mixture 90% N2-10% CO2 modeling the flue gas. With 5A zeolite having given the best performances, we tested it with various operating conditions including one with nitrogen purge during desorption. This one showed a good compromise between CO2 capture rate, purity of the desorbate, volumetric productivity, and specific-heat consumption. We obtained a volumetric productivity of 37 kgCO2/m3ads‚h and a specific-heat consumption of 6 MJ/kgCO2 at our laboratory scale and 4.5 MJ/kgCO2 for the adiabatic estimate (in the same order of magnitude as those obtained industrially with the reference MEA amine process). These results are promising because our process is not optimized yet and the scale-up on an industrial version involves a reduction in specific-heat consumption. 1. Introduction To deal with the greenhouse gases emission increase, CO2 capture and storage is one of the possible ways to lower emissions due to fossil fuel use for electricity production.1,2 Different technologies for CO2 capture are currently under development: membrane separation, physical/chemical absorption, and adsorption. Two options are considered: postcombustion CO2 capture (CO2 is removed from the flue gases) and pre-combustion capture (CO2 is removed before the combustion chamber, which will be the case for the power plant using a gasification process for example). If the separation is easier in the second case, as the CO2 molar fraction is high (typically about 30%), post-combustion capture is the main way to lower the emission of existing power plants and future power plants where CO2 is produced during the combustion. Furthermore, the post-combustion CO2 capture process can be used in the cement or steel industry to lower their furnace CO2 emissions. The major difficulty of post-combustion capture is to produce a highly concentrated CO2 stream matching the purity requirement for transportation while the CO2 is diluted in the flue gas: between 4% for the natural gas combined cycle (NGCC) and 14% for pulverized coal (PC). Of course, the energy consumption has to be low while keeping high CO2 recovery. Chemical absorption (conventional MEA) is todays reference process; it requires an energy cost ranging from 4.2 to 4.8 MJ/kgCO2.3 Adsorption processes were first not really considered as competitive for CO2 capture2 when compared to absorption. However, some recent studies have highlighted the potential of the pressure/vacuum swing adsorption (PSA/VSA) for CO2 capture.4,5,6 Park et al.4 found energy consumption in the range of 0.09-1.1 MJ/kgCO2 (0.025-0.3 kWh/kgCO2) for purity ranging from 50 to 70% and recovery between 30 and 90% after one-stage PSA (99% purity can be achieved by means of a two-stage PSA). A purity of 99% is obtained in Pilarczyk et al. ,7 with an energy consumption between 2.9 and 10.6 MJ/ kgCO2 (recovery between 53 and 72%). By simulating a FVPSA * To whom correspondence should be addressed. Tel.: +33 1.58.80.85.49. Fax: +33 1.40.27.20.47. E-mail: [email protected].

process, Ko et al.6 claim that a specific consumption as low as 0.14 MJ/kgCO2 can be obtained. In the low part of the previous range values, the energy consumptions are lower than those of amine processes, but the consumed energy is of the mechanical work type, whereas heat is directly used in chemical absorption. For these reasons, we have chosen to evaluate the potential of post-combustion capture by temperature swing adsorption (TSA). Contrarily to PSA/VSA, TSA can be directly heat driven. However, the main drawbacks of TSA are its low productivity, which results in a large adsorbent amount and desorbate dilution because of regeneration by hot gas purging.8,9 To avoid these drawbacks, an indirect TSA process developed in our laboratory is used.10,11 The originality of this process comes from the indirect heating during the regeneration step using an internal heat exchanger. Heating is performed in a two-phase heat transfer mode, namely condensation of steam. This allows us to reach high heat transfer coefficients, which reduces the regeneration time. During the adsorption step, the adsorber is cooled by water circulation that allows us to remove the adsorption heat and limits the impact of the inlet gas temperature. Then, the adsorbent capacity is kept maximal.12 The aim of the present work is the evaluation of the performances of this TSA process for CO2 capture with 13X and 5A zeolites. For this purpose, experiments will be carried out with different operating conditions, and the performances will be evaluated on the basis of CO2 capture rate, purity of the desorbate, productivity and specific-heat consumption. Of these results, the performances of the process on an industrial scale will then be evaluated. 2. Experimental Set-Up 2.1. Adsorber Configuration and Instrumentation. The experimental apparatus (Figure 1) is exhaustively described in Bonjour et al.10 so that only the main parameters are presented herein. The adsorber is a 1 m high heat exchanger with 12 external fins (Figure 1). It is made of concentric tubes (innertube outer diameter, 19 mm; outer-tube inner diameter, 70 mm), to form an annulus (internal volume, 3.65 dm3) filled with an adsorbent. During desorption, steam (min 130 °C, max 180 °C) condenses inside the inner tube, whereas tap water (12-15 °C)

10.1021/ie071012x CCC: $40.75 © 2008 American Chemical Society Published on Web 12/11/2007

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Figure 1. Schematic experimental setup and cross-section view of the adsorber.

Figure 2. CO2 adsorption equilibrium data on 13X and 5A zeolites.

Figure 3. N2 adsorption equilibrium data on 13X and 5A zeolites.

is used during the adsorption step for cooling. During the desorption phase, the condensed water is collected in a tank equipped with a level indicator, which allows us to measure the steam consumption. A differential manometer is used to measure the pressure drop through the porous bed. The gas mixture composition is determined at the outlet of the column by gas chromatography (TCD detector). The inlet composition is fixed by the ratio of the flow rates of each gas, which is adjusted by mass-flow controllers (uncertainty of 1%). Several thermocouples with an accuracy of about (1 K are placed at various horizontal and vertical positions. Hence, the temperature measurements combined with the steam consumption allow us to calculate the distribution of the energy consumption (adsorbent, losses to the ambient, etc.). 2.2. Adsorbents. The 13X and 5A zeolites are the commercially available adsorbents mostly studied for CO2 capture (particularly the 13X one). Hence, we have chosen to evaluate our process with both of them. The used 13X and 5A zeolites were supplied by AXENS (IFP Group Technologies) and are small spherical beads (average diameter of 2 mm) with a bulk density of 0.63 kg/l (2.31 kg in the adsorber) and 0.73 kg/l (2.66 kg in the adsorber), respectively. Typical flue gas consists of nitrogen, carbon dioxide, oxygen, water, and minor impurities such as SOx and NOx. Impurities and moreover water are known to have deleterious effect on 13X and 5A zeolites adsorption capacities so that they are supposed to be removed by a pretreatment process. For both zeolites, nitrogen has a stronger affinity than oxygen13,14 so that

dry exhaust gases could be well simulated with a binary N2/ CO2 mixture. The inlet molar concentration of CO2 is fixed at 10%, which is between 4% and 14%, respectively the rate of CO2 in the flue gas of power stations NGCC and PC.15 The experimental data of CO2 adsorption isotherms on 13X and 5A zeolites, coming from various authors (13X, Cavenati et al.;16 Harlick and Tezel,17 5A; Mulltoh & Finn;18 Pakseresht et al.19) were compared, and these data agreed well with each other, so that only experimental data from Cavenati et al.16 and Mulloth et al.18 are drawn in Figure 2. For the N2 adsorption isotherms, the experimental data plotted in Figure 3 are obtained from Cavenati et al.16 for 13X zeolite and Vereist and Baron20 for 5A zeolite. Figures 2 and 3 show that both adsorbents have a very high selectivity of CO2 over N2. The CO2/N2 ideal selectivity for the 13X and 5A zeolites is estimated from pure gas adsorption data:

R)

(n/C)CO2 (n/C)N2

(1)

In TSA operating conditions (25 °C/1 bar), the selectivity is high for both adsorbents: 96 (13X) and 90 (5A). On the other hand, with PSA conditions (25 °C/3 bar) (Gomes et al., 2002)21 and VSA conditions (25 °C/1.3 bar) (Chaffee et al., 2007),22 the ideal selectivity for both zeolites is around 48 and 81, respectively, for PSA and VSA. These ideal selectivities

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Figure 4. CO2 breakthrough curves for the 13X and 5A zeolites.

highlight the interest to run VSA or TSA processes instead of the classical PSA process. 3. Adsorption/Desorption Experiments

Figure 5. Temperature evolutions during breakthrough at various adsorber locations for the 13X and 5A zeolites. Table 1. Purge-Gas Configuration for Noncyclic Desorption

Preliminary experiments’ results are presented to determine the necessary parameters to decide on the time/criterion configuration for the cyclic experiments. 3.1. Breakthrough Experiments. The noncyclic experiments are carried out with a total volume flow rate of 20 Ndm3/min, a CO2 inlet molar fraction of 10%, and a cooling water of about 12-15 °C. The adsorption step is stopped at bed saturation. The breakthrough curves are reported in Figure 4. From Figure 4, it can be seen that with the 5A zeolite the first CO2 detection is later (+4 min) and the stoichiometric time (time at which the outlet CO2 concentration is half of that of the feed) is reached 13 min after that of 13X zeolite. By the integration of the curves, the total amounts of CO2 processed can be evaluated at about 349 g for the 13X zeolite and 396 g for the 5A one (+13.5% in mass), corresponding respectively to mass capacities of 0.151 and 0.149 kgCO2/kgads (same order of magnitude). These mass capacities are in accordance with the values obtained from the isotherms’ equations (Toth model) deduced from the experimental data of CO2 adsorption isotherms15,17 of section 2.2 (0.149 and 0.129 kgCO2/kgads, respectively, for 13X and 5A zeolites). These high capacities are obtained thanks to the indirect cooling of the bed during adsorption. We have also considered the volumetric capacities because the necessary volume of the adsorbent is important data for the process design. The volumetric capacities are 95 and 108 kgCO2/m3ads (+14%), respectively, for the 13X and 5A zeolites. From all of these results, 5A zeolite appears to be more efficient than 13X for CO2 adsorption in our case. Figure 5 shows the temperature evolutions at different bed locations during the adsorption step for the two zeolites. For 13X, the rises in temperature due to CO2 adsorption occur before the 5A ones, confirming the shorter time of the 13X breakthrough. As it can be seen in Figure 5, at all the locations the temperature overshoots of the 13X are higher than those of the 5A (a difference decreasing from 20 (P1) to 6 K (P5)), but in both cases the temperatures reach 90% of their initial value after 30 min. So, the cooling is not a limiting factor for the temperature evolutions of the zeolites. With this internal heat exchanger, heat is removed from the adsorbent and gas; that is why all of the temperatures at the

purge-gas configuration

13X(a)

5A(1a)

5A(2a)

5A(3a)

N2 purge flow rate (Ndm3/min) preheating regeneration

0 2

0 2

noa 1

noa 0.2

a No preheating; desorption starts immediately with the regeneration (N2 purge gas).

different locations are close to 15 °C at the end of adsorption step, which preserves the bed capacity. 3.2. Desorption Experiments. The desorption could be divided into two steps: preheating and regeneration. During the desorption, the steam circulating in the internal heat-exchanger is at 150 °C. For preheating, no purge gas is used so that the flow rate is only due to the thermal expansion coupled with CO2 desorption. When the flow rate of the desorbate becomes lower than 1.2 Ndm3/min, a nitrogen purge gas is admitted to complete the regeneration. The different purge gas configurations tested are listed in Table 1. The desorption step is stopped when the bed is clean. It should be noted that the 5A(1a) experiment is realized with the same desorption operating conditions as those of the 13X to obtain a comparison of both. With 5A zeolite showing better performances, we only tested different desorption operating conditions for this zeolite. Figure 6 shows the history of the averaged outlet gas composition phase and the evolution of the temperature, at 0.5 m from the top of the column, during the regeneration for the four different desorptions. We can see in Figure 6 that the four temperature evolutions are similar. A slight slope change can be observed for 13X(a) and 5A(1a) during the first minutes after the introduction of 2 Ndm3/min of gas purge. We can conclude that the operating conditions have only a slight influence on the temperature evolutions during desorption. For the 13X(a) and 5A(1a) experiments, the history of the averaged outlet CO2 molar concentration in the desorbate is around 99% during the preheating step. Indeed, even if no nitrogen purge is used, there is a part of adsorbed and interstitial nitrogen. When desorption starts directly with the regeneration step, the desorbate purity reaches only 91 and 98% after the first 6 min for 5A(2a) and 5A(3a), respectively. These results highlight that the use of preheating results in higher CO2 purity.

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Ind. Eng. Chem. Res., Vol. 47, No. 1, 2008 Table 2. Operating Conditions for the Cyclic Experiments operating conditions

13X(b) 5A(1b) 5A(2b)

5A(3b)

adsorption total flow rate (Ndm3/min) CO2 inlet molar fraction, C0 (%) cooling water temperature (°C) criterion of the end of adsorption adsorption time (min)

20 10 12-15 0.5 C0 49

20 10 12-15 49 min 49

20 10 12-15 0.5 C0 74

20 10 12-15 1st increase 50

desorption nitrogen purge flow rate (Ndm3/min) preheating regeneration Temperature of steam water (°C) Criterion of the end of desorption

0 nob 150 16 min

0 nob 150 16 min

0 nob 150 22 min

noa 0.2 150 15 min

a No preheating step; desorption starts directly with the regeneration step with nitrogen purge. b No regeneration step; desorption stops after the preheating.

Figure 6. History of the averaged outlet CO2 molar fraction and temperature evolution during the four types desorptions.

involving good purity (>97%) and productivity (a purge of 0.2 Ndm3/min giving better results than 1 Ndm3/min). 4. Cyclic Experiments

Figure 7. Cumulated desorbed CO2 mass versus time evolution.

Although the achieved purities with the proposed adsorption process are lower than those obtained with MEA (>99%),2 these purities are high enough to comply with the constraints of CO2 transport and storage.2 Because of the lower flow rates of the purge gas, the purity of the desorbate decreases more slowly in the 5A(2a) and 5A(3a) desorptions than in the 13X(a) and 5A(1a) cases. Then, as it can be seen in Figure 6, the history of the averaged outlet molar concentration stays much longer above 90% for the 5A(3a) desorption (92 min) than the 13X(a) and 5A(1a) ones (respectively 24 and 30 min). In Figure 7, the cumulated desorbed CO2 mass curves are plotted as a function of time for the four experiments. With the same operating parameters (13X(a) and 5A(1a)), both zeolites show similar times to desorbe 50% of CO2 (around 18.5 min), and at this time the specific-energy consumptions are in the same range (5 MJ/kgCO2). When we compare the results obtained with 5A zeolite, it can be noted that the preheating disadvantages desorption in terms of desorption rate: the times to desorb 50% of CO2 are 12.1 and 16.1 min for 5A(2a) and 5A(3a), respectively. However, the specific consumptions are approximately the same (5 MJ/kgCO2). From all of these results, two options appear for the desorption operating conditions: no gas purge used involving a very good purity (>99%) but a low productivity; or a gas purge use

4.1. Experimental Results and Performances. The full cyclic process includes two main steps: adsorption and desorption (could be divided in two steps, preheating and/or regeneration). The different tested operating conditions are listed in Table 2. It has to be noticed that the criteria of the end of the adsorption are chosen arbitrarily: 0.5 C0 to get a small LUB (length of unused bed), so that the bed is almost totally saturated (an important breakthrough is possible because there is no need to capture 100% of the CO2); first increase of detected CO2 (not the first appearance of CO2) because during cyclic experiments the outlet CO2 molar fraction does not reach zero. For clearness, the outlet gas composition during successive cycles (adsorption/desorption) and the temperature evolution are only plotted for the 5A(2b) experiment in Figure 8. It has to be noticed that the shapes of the curves for the other experiments are similar. For all four experiments, the outlet CO2 molar fraction and the temperature exhibit different evolutions when comparing the first cycle to the following ones. This is due to a different initial state of the adsorber at the beginning step. Indeed, for the first adsorption the bed is at low temperature (15 °C) and free of CO2. Consequently, the adsorption capacity is maximum so that the CO2 is adsorbed quickly in a great amount, resulting in a high heat generation (temperature overshoot of 30 K). On the contrary, for the following adsorptions, the adsorber is initially at high-temperature (above 120 °C), and the adsorbent is partially saturated (regeneration stopped before complete cleaning). Then, when the bed is cooled down, less CO2 is adsorbed, resulting in a low-temperature overshoot (10 K). Another consequence is that the outlet CO2 molar fraction does not reach zero during cyclic adsorption. This is of minor importance as the objective of this kind of process is not to recover 100% of the CO2. Four comparison criteria are considered: CO2 capture rate (%CO2), recovered CO2 average purity (%CO2), volumetric productivity (kgCO2/m3ads‚h) and specific-heat consumption (MJ/ kgCO2). All of these results are shown in Figure 9. With the same operating conditions, 5A shows better results than 13X for three selected criteria: CO2 capture rate (+14.5%), volumetric productivity (+22%), and specific-heat consumption (-19%). These cyclic results confirm the noncyclic ones: it is much more interesting to use 5A zeolite than 13X. That is why we carried out other tests in cycle only on the 5A to determine more-optimal operating conditions.

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Figure 8. Outlet CO2 molar fraction and temperature evolution during the 5A(2b) cycle.

Figure 9. Comparison of the performances of the four cycles.

Between the 5A(1b) and 5A(2b) conditions, the last one seems to be more efficient. Indeed, if purities and productivities are on the same order of magnitude; for 5A(2b), the CO2 capture rate is higher by 9% and the specific consumption is lower by 10% than those of 5A(1b). With a nitrogen purge (0.2 Ndm3/ min) in 5A(3b), the volume productivity increases by 13.5% and the specific consumption decreases by 15.5% compared to those of 5A(1b), but the CO2 recovered purity is only 94%. [Contrarely to 5(1b) and 5(2b) cases, the achieved purity is lower but stay close to what could be the acceptable value for compression and storage (95%)]. However, this desorbate purity is slightly lower than those of absorption by amine (>99%)2 and is largely acceptable for the later stages as storage (>95%).2 It has to be noticed that the purities of the desorbate found experimentally are on the same order of magnitude (94, 97, 98, and 94%, respectively, for 13X(b), 5A(1b), 5A(2b), and 5A(3b)). 4.2. Energy Requirement Distribution. Energy requirement calculations were only carried out from the experiments presenting the best results, that is, 5A(2b) and 5A(3b). The energy required for the heating of the adsorbent and the metal parts of

the adsorber (steel) is calculated using the specific heat of each element and the corresponding temperature variation. The heat used for desorption is directly linked to the amount of operated CO2 (adsorbed N2 amount is neglected) and the isosteric heat of adsorption for 5A zeolite is 1023 kJ/kg.18 On the other hand, the heat losses are deduced from the difference between total consumed energy (calculated from the mass of condensed steam), the heat of desorption, and the energy necessary for the heating of the various elements of the adsorber. The mass of the metal parts is 6.8 kg (inner tube, 1.2 kg; fins, 2.3 kg; jacket, 3.3 kg) that makes a total mass (adsorber + adsorbent 5A) of 9.46 kg. Results are presented in Table 3. From the results of Table 3, we note that the energetic performances of our process are slightly better with the 5A(3b) experimental protocol (6.12 MJ/kgCO2) than with the 5A(2b) (6.46 MJ/kgCO2). But this specific-heat consumption is higher than that of absorption by an amine in the industrial version (4.2-4.8 MJ/kgCO2 according to Aroonwilas and Veawab3). However, it should be considered that in the case of our adsorber at laboratory scale, the heat losses are important (32 and 25%, respectively for 5A(2b) and 5A(3b)), which is not

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Table 3. Energy Requirement of the Process

Table 4. Scale-Up to the Industrial Scale 5A(2b)

5A(3b)

Adsorption adsorption time (min) amount of processed CO2 (g)

74 224

50 136

N2 purge flow rate (Ndm3/min) preheating regeneration Desorption time (min)

0 no 22

no 0,2 15

Energy for heating (MJ) total energy requirement adsorbent inner tube fins jacket desorption heat heat losses specific-heat consumption (MJ/kgCO2) specific-heat consumption (adiabatic version) (MJ/kgCO2)

1.102 0.239 0.082 0.155 0.102 0.174 0.349 6.46 4.41

0.894 0.212 0.080 0.151 0.081 0.150 0.222 6.12 4.60

Desorption

representative of an industrial-size adsorber as mentioned by Bonjour et al.23 Consequently, it is necessary to consider the specific consumption of our process in an industrial scale in order to obtain a fairer comparison with the processes of absorption by an amine. At first approximation, the process is supposed to be adiabatic and the total energy consumption without heat loss is then evaluated: sum of the necessary energies for the heating of the adsorber and adsorbent, and the desorption heat. These estimations of the adiabatic version, given in Table 3, highlight the important reduction in specific-energy consumption (4.41 and 4.60 MJ/kgCO2 against 6.46 and 6.12 MJ/kgCO2, respectively, for 5A(2b) and 5A(3b)). Furthermore, these results are in the range of the conventional amine absorption process. In the case of the adiabatic version, the 5A(2b) conditions become more interesting than the 5A(3b) conditions. This is explained by the total required energy in 5A(2b) that decreases more than that in 5A(3b) when the heat losses are omitted. Indeed, these are more important for 5A(2b) due to the longer duration of the desorption step (cf. Table 3): 32% and 25% of heat losses for 5A(2b) and 5A(3b), respectively. These results show that there is a good potential of energy saving as the configuration of our process is not optimized yet (number of beds, cycles duration, etc.), and at industrial scale the adsorbent/ metal part ratio is more favorable, involving a reduction in specific-energy consumption.23 4.3. Scale-Up for a Power Plant with CO2 Capture. A scale-up of our indirect TSA process is interesting in order to estimate its size in the industrial case and determine if it is feasible. It has to be noticed that, in this part, the scale-up is realized in first approximation with some assumptions concerning the heat transfer: the thermal performances are supposed to be the same between the laboratory and industrial scales. For the scale-up it is necessary to have an example of power plant compatible with our experimental conditions (10% of CO2). The biomass power plants with CO2 capture, resulting in negative emissions of CO224, 25, 26, 27 are interesting. The study is carried out from a typical biomass power plant of 30 MWel with an emission of 0.2 MtCO2/an (22.8 tCO2/h).2 We also consider large beds of 5 m diameter and 10 m height (196 m3) and a parallel adsorption/desorption run. The results of the study of the scale-up are summarized in Table 4. The results of Table 4 show the important difference concerning the dimension of the process according to the zeolite used, indeed the process needs five columns with the 13X zeolite and four columns with the 5A one, in the same operating

volumetric productivity (kgCO2/m3ads‚h) volume of adsorbent (m3) mass of adsorbent (t) number of columns for the adsorption phase

13X(b)

5A(1b)

5A(2b)

5A(3b)

25 911 574 5

32 716 523 4

30 761 556 4

37 616 450 3

conditions. Moreover, with the 5A zeolite, the number of necessary columns decreases from four to three with the mostfavorable operating conditions. At first approximation of the configuration of our process, we consider one desorption column by adsorption column, which gives a total volumetric productivity of 18.5 kgCO2/m3ads‚h and six columns for 5A(3b) conditions. The installation of this process is realistic, taking into account the number of necessary columns of adsorbent varying between six and eight (i.e., from 1232 to 1522 m3). On the other hand, according to Tobiesen and Svendsen,28 for the CO2 capture pilot plant (amine process) the total volumetric productivity is around 39 kgCO2/m3abs‚h, involving 585 m3 of MEA. But it is important to note that complementary works on cyclic configurations, adsorbent choice, and energy efficiency can reduce the size of the installation. 5. Conclusion In the present work, the performances of a TSA process with indirect heating and cooling by means of an internal heatexchanger were investigated experimentally on 13X and 5A zeolites for post-combustion CO2 capture. The flue gas was simulated by a N2/CO2 mixture, and for all tested operating conditions, very good separation is achieved as CO2 is recovered at very high purity (g94%). Zeolite 5A shows better results than the 13X for three criteria: CO2 capture rate (+18%), volumetric productivity (+23%), and specific-heat consumption (-19%). From different cyclic experiments on the 5A zeolite, it was found that purge-gas use (0.2 Ndm3/min) gives better volumetric productivity and specific-heat consumption than when no purge is used: +14% and -15.5% (energy saving) respectively. Furthermore, the purity of the desorbate decreases from about 98 to 94%. These results are in the same order of magnitude as those obtained with absorption by an amine concerning the purity (about 99%) and slightly lower for the CO2 capture rate (between 75 and 85%). On the other hand, the specific-heat consumptions (6.12 and 6.46 MJ/kgCO2 for the most-favorable conditions) are higher than that of the absorption with an amine (4.2 to 4.8 MJ/kgCO2). However, it should be considered that in the case of our adsorber at a laboratory scale, the heat losses are high, which are not representative of an industrial-sized adsorber.23 An estimate of the adiabatic version shows that the specific-energy consumption (4.41 and 4.60 MJ/kgCO2 for the most-favorable conditions) is then on the same order of magnitude as that of absorption with amine. These results are promising because, with a scale-up, the adsorbent/metal ratio is then more favorable, involving a reduction of the specificheat consumptions.23 Moreover, it is important to note that the performance of our TSA process can still be improved by complementary studies on the choice of the adsorbent, the configuration of the cycles, and the energy efficiency. To go in this direction, further work will consist in: • experimentally testing other operating conditions • developing a simulation tool to run a complete parametric study

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• integrating the capture system in a power-plant process, taking into account the needed intermediate steps (flue gas water removal for example) Nomenclature C ) gas-phase concentration (kg/m3) n ) adsorbed phase concentration (mol/kgads) Greek Letters R ) ideal selectivity Subscripts abs ) absorbent ads ) adsorbent el ) electric Literature Cited (1) Williams, R. H. Toward Zero Emissions from Coal in China. In China Clean Energy Forum; Beijing, 2001. (2) IPCC. Special Report on Carbon Dioxide Capture and Storage, Prepared by Working Group III of the Intergovernmental Panel on Climate Change, Metz, B., Davidson, O., de Coninck, H., Loos, M., Meyer, L. A., Eds.; Cambridge University Press: Cambridge, U.K. and New York, U.S.A., 2005. (3) Aroonwilas, A.; Veawab, A. In Cost Structure and Performance of CO2 Capture Unit Using Split-Stream Cycle, Proceedings of the 8th International Conference on Greenhouse Gas Control Technologies (GHGT8), Trondheim, Norway, 2006. (4) Park, J. H.; Beum, H. T.; Kim, J. N.; Cho, S. H. Numerical Analysis on the Power Consumption of the PSA Process for Recovering CO2 from Flue Gas. Ind. Eng. Chem. Res. 2002, 41, 4122. (5) Ko, D.; Siriwardane, R.; Biegler, L. T. Optimisation of Pressure Swing Adsorption Using Zeolite 13X for CO2 Sequestration. Ind. Eng. Chem. Res. 2003, 42, 339. (6) Ko, D.; Siriwardane, R.; Biegler, L. T. Optimisation of Pressure Swing Adsorption and Fractionated Vacuum Pressure Swing Adsorption Processes for CO2 Capture. Ind. Eng. Chem. Res. 2005, 44, 8084. (7) Pilarczyk, E.; Schro¨ter, H. J. In New PSA-Processes with Carbon Molecular SieVed for RecoVery of Carbon Dioxide and Methane, Proceedings of the Gas Separation Technology conference, Antwerpen, Belgium, 1989. (8) Ruthven, D. M. Principles of Adsorption and Adsorption Processes; John Wiley & Sons: New York, 1984. (9) Yang, R. T. Gas Separation by Adsorption Processes; Imperial College Press: London, 1997. (10) Bonjour, J.; Chalfen, J. B.; Meunier, F. Temperature Swing Adsorption Process with Indirect Cooling and Heating. Ind. Eng. Chem. Res. 2002, 41, 5802. (11) Clausse, M.; Bonjour, J.; Meunier, F. Influence of the Presence of CO2 in the Feed of an Indirect Heating TSA Process for VOC Removal. Adsorption 2003, 9, 77. (12) Clausse, M.; Bonjour, J.; Meunier, F. Adsorption of Gas Mixtures in TSA Adsorbers under Various Heat Removal Conditions. Chem. Eng. Sci. 2004, 59, 3657.

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ReceiVed for reView July 25, 2007 ReVised manuscript receiVed September 26, 2007 Accepted October 1, 2007 IE071012X