Energy & Fuels 2006, 20, 721-726
721
Material Balance and Energy Consumption for CO2 Recovery from Moist Flue Gas Employing K2CO3-on-Activated Carbon and Its Evaluation for Practical Adaptation Naoya Shigemoto* and Tetsu Yanagihara Chemical Technology Department, Shikoku Research Institute, Inc., 2109 Yashima-nishimach, Takamatsu, Kagawa, 761-0192, Japan
Shigeru Sugiyama and Hiromu Hayashi Department of Chemical Science and Technology, UniVersity of Tokushima, 2-1 Minami-josanjima, Tokushima, 770-8506, Japan ReceiVed August 12, 2005. ReVised Manuscript ReceiVed January 18, 2006
Potassium carbonate supported on an activated carbon has been proposed as an efficient sorbent to recover CO2 from moist flue gases. As a characteristic of the present CO2 sorption process, which can be described as K2CO3‚1.5H2O + CO2 ) 2KHCO3 + 0.5H2O, moisture in the feed gases had no influence on the CO2 sorption. By the temperature-swing operation of a fixed-bed, the CO2 recovery was achieved as follows: carbon dioxide in moist flue gases at around 363 K was sorbed by the K2CO3 sorbent, followed by steam flushing at 433 K to release the CO2, and then cooling the sorbent for the next CO2 sorption. In the present study employing a bench-scale apparatus, the material (CO2 and H2O) balances, together with those of heat during each step, were measured to elucidate the CO2 sorption/release and the cooling behaviors. To evaluate the practical adaptability of this process, the heat consumption for the CO2 recovery on a commercial-scale was estimated. When compared with that for other processes such as the conventional amine process, it provided a remarkable energy-conservative effect. The cost for the CO2 recovery by K2CO3-on-activated carbon is also discussed.
1. Introduction
K2CO3‚1.5H2O + CO2 ) 2KHCO3 + 0.5H2O
A variety of methods, such as chemical absorption, physical adsorption, and membrane separation, have been used to recover CO2 from the flue gases of large combustion plants. However, moisture in the flue gases adversely affects the adsorption capacity of CO2 and should be removed before the gases are supplied to the conventional adsorption process. A chemical absorption method by an aqueous potassium carbonate as employed in the Benfield process1,2 has no concern with the moisture in flue gases, but the method consumes a sizable amount of primary energy for heating a large volume of solvent water to release the entrapped CO2. Thus, to minimize the energy consumption for the CO2 recovery, potassium carbonate employed in the chemical absorption process was supported on porous materials such as an activated carbon (AC). In previous studies3,4 employing a lab-scale column packed with K2CO3on-AC, it was revealed that moisture in the feed gases showed no ability to sorb CO2 while the hydrate of K2CO3 was held in the pores of the activated carbon to reversibly sorb/release CO2 by the following reaction:5
However, the lab-scale experiment was too small to evaluate the energy consumption for the CO2 recovery. Although K2CO3-on-AC could adapt to moist flue gases and consume steam in the CO2-releasing step, the behavior of water (moisture and steam) in the CO2 recovery process has not been investigated. Material balances for both the CO2 and water through the CO2 recovery process were of significant importance to realize the K2CO3-on-AC process. Thus, in the present study, a bench-scale experiment of the CO2 sorption/release was carried out to examine the material balances of CO2 and H2O together with an energy balance in the CO2 recovery process with K2CO3-on-AC. Moist CO2 was supplied to K2CO3-on-AC packed in a bench-scale column (54.5 mm i.d. × 800 mm length) to sorb CO2, followed by steam flushing to release the entrapped CO2. The released gas was then cooled through a heat exchanger to condense the steam and afford CO2 in high purity. In the CO2 sorption/release and the sorbent cooling steps, a material balance and behavior of the CO2 and water were elucidated. An accumulated water amount on the sorbent during the steam-flushing step, which is obtained from the water balance, especially offers energy consumption for the CO2 release using the steam latent-heat supplement. The specific heat of K2CO3-on-AC and the heat of the CO2 sorption reaction were also measured to estimate the heat for warming the sorbent and the decomposition heat of the formed KHCO3, respectively. By assuming a commercialscale CO2 recovery plant, the energy consumption was estimated and compared with that for other chemical absorption processes,
* Corresponding author. Phone: +81-87-844-9216. Fax: +81-87-8449236. E-mail:
[email protected]. (1) Bartoo, R. K. Chem. Eng. Prog. 1984, 80, 35. (2) Benson, H. E.; Field, J. H. U.S. Patent 2,886,405, 1959. (3) Hayashi, H.; Hirano, S.; Shigemoto, N.; Yamada, S. Nippon Kagaku Kaishi 1995, 1006. (4) Hirano, S.; Shigemoto, N.; Yamada, S.; Hayashi, H. Bull. Chem. Soc. Jpn. 1995, 68, 1030. (5) Hayashi, H.; Taniuchi, J.; Furuyashiki, N.; Sugiyama, S.; Hirano, S.; Shigemoto, N.; Nonaka, T. Ind. Eng. Chem. Res. 1998, 37, 185.
10.1021/ef058027x CCC: $33.50 © 2006 American Chemical Society Published on Web 02/17/2006
(1)
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Figure 2. CO2 breakthrough curve. Feed gas: 13% CO2/N2 + 10% H2O, 4.8 L/min.
Figure 1. Schematic flow diagram of CO2 recovery apparatus in a bench-scale apparatus. Column: 54.5 mm i.d. × 800 mm length. MF: Mass flow meter.
thus demonstrating the practical applicability of the K2CO3-onactivated carbon process. The cost estimation for the CO2 recovery is also discussed. 2. Experimental Section 2.1. Preparation of K2CO3-on-Activated Carbon. An aqueous solution (800 mL) of 26 wt % K2CO3 with a small amount of a surfactant (alkylbenzenesulfonic acid type, Lipon PS-260, Lion Co., Tokyo) was prepared (weight ratio of the surfactant/K2CO3 ) 1/35). The K2CO3 solution was added to 1.00 kg of a dry activated carbon extrudate (WH2X type, 4-8 mesh, Takeda Chemical Industry, Co., Osaka, Japan), and then the mixture was dried in a rotary evaporator under reduced pressure at 333 K. The obtained solid particles were dried again in a heated oven at 378 K. These sorbent particles contained 26.4% (1.92 mmol/g dry sorbent) K2CO3 and 3.5% moisture. 2.2. Bench-Scale Apparatus and Its Operation. A schematic flow diagram of a bench-scale apparatus employed in the present study is shown in Figure 1. A column of 54.5 mm i.d. and 800 mm length (1.87 L) was packed with 1.24 kg of K2CO3-on-AC prepared as previously described. The column was covered with an annular thermal insulator made of glass wool to prevent heat loss through the outer surface of the column. The temperatures of the sorbent were measured by three thermocouples located at the lower position (inlet side), the center, and the upper position (outlet side) of the column as shown in Figure 1. A moist CO2 gas (10% H2O, 11.8% CO2), which was prepared by the addition of steam to dry 13.0% CO2/N2 gas, was supplied at the constant flow rate of 4.8 L/min from the bottom of the column at around 363 K. The flow rate and CO2 concentration at the column top were measured by a mass flow meter and an infrared CO2 analyzer (PG-200, Horiba Seisakusho, Tokyo, Japan), respectively. After the CO2 breakthrough, high-temperature steam (433 K, 0.618 MPa) was flushed through the sorbent from the column bottom at a constant flow rate (5.52 kg/h). The released gas from the column top was cooled through a heat exchanger to separate the water after condensation, and then the flow rate and CO2 concentration were measured. The hot column after the steam-flushing step was cooled under vacuum and then by flowing air at room temperature (293 K) to the initial temperature. To establish the steady state of the
sorbent, the following steps were successively repeated: CO2 sorption, 4.8 L/min × 10 min; steam (at 433 K) flushing, 5.52 kg steam/h × 1 min; vacuum cooling, 10 min; and air-flow cooling, 10 L/min × 5 min. After several cycles of successive operation, the bench-scale experimental data were collected. 2.3. Measurement of Sorbent Specific Heat and Heat of Reaction. The activated carbon extrudate impregnated with K2CO3 was crushed, sieved into particles of 1-2 mm in diameter, dried overnight in helium at 403 K, and cooled to room temperature. Employing a portion of the sample, the specific heat of K2CO3on-AC was measured by a differential scanning calorimeter (DSC7, Perkin-Elmer Co., Wellesley, MA), in which the sample was heated at 10 K/min under N2. The heat of reaction of CO2 with K2CO3/C was also measured by employing a portion of the same sample to a twins-type, conductive calorimeter (C80 type, SETARAM Co., France), where CO2 (>99.9%) was supplied to the sample at the flow rate of 10 mL/min while the heat change was documented to calculate the heat of the reaction.
3. Results 3.1. Carbon Dioxide Breakthrough Curve and Temperature Change of Sorbent during the CO2 Sorption Step. Figure 2 shows the CO2 breakthrough curve, where moist 13% CO2 gas was supplied to K2CO3-on-AC packed in the benchscale column. The concentration of CO2 at the column outlet indicated zero during the start to around 2000 s and then abruptly increased to attain the inlet CO2 concentration. Although after several repetitions of a series of CO2 sorption/releases and cooling steps, the initial temperatures at the lower, the center, and the upper positions were somewhat different from the same first 363 K, the temperatures at each position of the column abruptly increased after different time periods and then formed peaks as shown in Figure 3. It should be noted that the temperature peaks were observed in the order of the lower (inlet side), the center, and the upper positions (outlet side) of the column with the reaction time, presumably due to heat of the reaction of the impregnated K2CO3‚1.5H2O with the supplied CO2 as shown by eq 1. 3.2. Temperature Change and CO2 Release Flow Rate during the Steam Flushing Step. After the CO2 was sorbed, steam at 433 K was flushed through the sorbent to release the entrapped CO2. The temperature at the column center and the flow rate of the released gas after moisture separation were monitored as shown in Figure 4a and b, respectively. The CO2 concentration in the released gas was more than 99%, and thus
Material Balance and Energy Consumption for CO2 RecoVery
Figure 3. Temperature profile of column in the CO2 sorption step. Position: Inlet side (O), center (2), outlet side (0). Other conditions as in Figure 2.
Energy & Fuels, Vol. 20, No. 2, 2006 723
Figure 5. Transition of column temperature in the sorbent cooling step. Under vacuum 0-600 s. Air flow: 600-900 s, 10 L/min.
Figure 6. Effect of temperature on reaction heat of K2CO3-on-AC with CO2.
Figure 4. Temperature change (a) and released CO2 flow rate (b) in the steam flushing step. Steam flow rate: 5.52 kg/h.
the flow rate of the released gas almost agreed with that of the released CO2. By steam flushing, the temperature at the column center suddenly increased to attain a plateau at 433 K. The flow rate of the released CO2 abruptly increased with the column center temperature increasing and formed a sharp peak, followed by an abrupt decrease to zero. 3.3. Temperature Change during the Cooling Step. The hot sorbent after the steam flushing was cooled under vacuum and followed by flowing air as shown in Figure 5, where the sorbent temperature gradually decreased to 379 K and the final bed pressure was attained to 12.5 kPa in the vacuum cooling step, and then the sorbent temperature returned to the initial temperature for the next CO2 sorption in the succeeding airflow cooling step. 3.4. Temperature Dependences of Reaction Heat and of Sorbent-Specific Heat. The heat of the CO2 sorption at 25 °C was calculated to be -37.9 kJ/mol (exothermic) as the difference in the standard formation enthalpies for the components included in eq 1 as follows:6 K2CO3‚1.5H2O, -1609 kJ/mol; CO2(g), -393.7 kJ/mol; KHCO3, -959.8 kJ/mol; H2O(g), 242.0 kJ/mol.
The product of the heat of the reaction, 37.9 kJ/mol, at 298 K and the K2CO3 content of 1.92 mol/kg sorbent leads to the theoretical reaction heat per unit mass of the sorbent of 72.8 kJ/kg sorbent at 298 K. Figure 6 indicates the effect of temperature on the heat of the reaction, where the heat of the reaction increases in proportion to 1/T. From Figure 6, the heat of the reaction was extrapolated to 65.7 kJ/kg at 298 K, which corresponded well to the previously mentioned theoretical value of 72.8 kJ/kg sorbent. To estimate the heat for warming the sorbent during the steam flushing, the specific heats of the sorbent at different temperatures such as the CO2 sorption temperature and its release temperature were necessary. The temperature dependence of the K2CO3-on-AC on the specific heat was then experimentally measured as shown in Figure 7, in which the specific heat increased in proportion to the temperature. 4. Discussion 4.1. Carbon Dioxide Balance. The carbon dioxide balance in the overall CO2 recovery process is summarized in Table 1, where the CO2 amounts of supply, accumulation, and release (6) Chemical Handbook, Basic, 3rd ed.; Nippon Kagaku Kai, Maruzen: Tokyo, 1988; pp 306-308.
724 Energy & Fuels, Vol. 20, No. 2, 2006
Shigemoto et al. Table 3. Heat Balancea in the CO2 Recovery Process process
heat source
distribution
CO2 sorption
23.4 (heat of CO2 sorption)
steam flushing
179.8 (latent heat of steam)
21.5 (evaporation of released H2O) 4.0 (decomposition heat for KHCO3) 71.4 (temp rise of sorbent) 105.2 (temp rise of vessel)
column cooling
71.4 (temp rise of sorbent) 126.0 (temp rise of vessel)
a
Figure 7. Temperature dependence of specific heat for K2CO3-onAC. Table 1. Carbon Dioxide Balancea during the CO2 Recovery Process process (period)
supply
accumulation
release
CO2 sorption (4000 s) steam flushing (400 s) column cooling (900 s) overall
1.55 0 0 1.55
1.05 0.09 0 0
0.50 0.967 0.09 1.56
a
Mol/kg dry sorb. Table 2. Water Balancea during the CO2 Recovery Process process
initial
supply
accumulation
release
final
CO2 sorption steam flushing column cooling overall
1.95 1.40 6.17 9.52
1.23 27.07 0 28.30
-0.55 4.77 -4.21 0.01
1.78 22.30 4.21 28.29
1.40 6.17 1.95 9.52
a
Mol H2O/kg dry sorb.
are indicated as mol CO2 per unit mass of the sorbent. By the integration of the CO2 sorption rates, which were obtained from the CO2 breakthrough curves (Figure 2), the total amount of the sorbed CO2 was 1.05 mol CO2/kg sorbent as shown in Table 1. The total amount of the CO2 sorption corresponded to 54.7% of the stoichiometric amount of the CO2 sorbed against K2CO3 supported on the AC (1.92 mol/kg sorbent). The integration of the released CO2 flow rates during the steam flushing gave the total amount of the released CO2 of 0.967 mol CO2/kg sorbent (Table 1), which was 92.0% of the total amount of the sorbed CO2. During the sorbent cooling step, 0.09 mol CO2/kg sorbent was exhausted out of the column. The sum of the released CO2 amount by the steam flushing and the exhausted CO2 amount during the cooling step was 1.05 mol CO2. This sum was equal to the total CO2 sorption amount () the accumulation amount in the sorption step) in Table 1, suggesting no accumulation of CO2 in the sorbent during the cyclic operation of each step. 4.2. Water Balance. Table 2 summarizes the water balance in each step for the CO2 recovery process, where the initial and the final water contents of the sorbent, the water amount supplied to the sorbent (moisture in feed gas and flushing steam), and that condensed through a heat exchanger are indicated. The accumulated amount of water was calculated by considering the water balance (i.e., the incoming and outgoing of water during each step). During the CO2 sorption step, the water content of the sorbent decreased from the initial 1.95 to the final 1.40 mol H2O/kg sorbent for the CO2 sorption step, in which the difference was
168.3 (evaporation of H2O on AC)
kJ/kg dry sorb.
0.55 mol H2O/kg sorbent to discharge from the sorbent. The discharging water amount expected by eq 1 with 54.7% of the stoichiometry, which was the same as that for the CO2 sorption as mentioned above, was 0.525 mol H2O/kg sorbent and agreed well with the measured value of 0.55 mol H2O/kg sorbent. This agreement in the calculated and measured amounts of the discharge water strongly supports eq 1 as the CO2 sorption reaction. It was suggested that the CO2 sorption reaction consumed the hydrated water in K2CO3‚1.5H2O, rather than the moisture in the feed gases which occurred even in the dry feed gases. The reason the discharged water could not be adsorbed on the AC was that the relative humidity of the feed gas at the CO2 sorption temperature (363 K) was low enough (14.5%) to evaporate the discharged water and not reach the saturated steam pressure. By the steam flushing for the CO2 release, the water content of the sorbent increased from the initial 1.40 to the final 6.17 mol H2O/kg sorbent, producing 4.77 mol H2O/kg sorbent of accumulated water at the sorbent. A portion of the accumulated water, 0.525 mol H2O/kg sorbent, should be required to regenerate K2CO3‚1.5H2O from KHCO3 as described by eq 1, but in reverse. The residual amount of the accumulated water, 4.25 mol H2O/kg sorbent, should then be allotted to the physical adsorption on the support AC. During the sorbent cooling step, the water content of the sorbent decreased from the initial 6.17 to the final 1.95 mol H2O/kg sorbent, in which the difference in the water content, 4.21 mol H2O/kg sorbent, should be exhausted out of the column. 4.3. Heat Balance. Table 3 summarizes the heat balance during each step for the CO2 recovery process, which was estimated on the basis of a material balance, especially for the water balance mentioned above, and the heat data, such as the heat of the reaction, the latent heat of water evaporation, and the specific heat of the sorbent. The reaction heat when the CO2 sorption reached 54.7% of stoichiometric should be 22.0 kJ/kg sorbent at the CO2 sorption temperature because the heat of the CO2 sorption at 363 K was expected to be 40 kJ/kg sorbent from Figure 6. On the other hand, the multiplication of the risen temperature (26 K temperature peak at the column center in Figure 3) and the specific heat of the sorbent at 363 K (0.90 kJ‚kg-1‚K-1) from Figure 7 gave 23.4 kJ/kg sorbent as the heat for warming the sorbent. The sorbent warming heat agreed well with the theoretical heat of the CO2 sorption at 363 K of 22.0 kJ/kg sorbent mentioned above. According to the actual CO2 sorption, 0.5 mol of the hydrated water of K2CO3‚1.5H2O, which corresponded to 0.525 mol/kg sorbent, should be released from the sorbent as shown in Table 2. The evaporation heat for the released water was calculated to be 21.5 kJ/kg sorbent, by employing the latent heat of water, 2280 kJ/kg H2O at 363 K. The coincidence of
Material Balance and Energy Consumption for CO2 RecoVery
the heat of CO2 sorption, 23.4 kJ/kg sorbent, and the evaporation heat for the released water, 21.5 kJ/kg sorbent, well explained that the initial and the final temperatures of the sorbent in the CO2 sorption step were similar, though the temperature peak due to the heat of the reaction appeared during the CO2 sorption step, as shown in Figure 3. The flushing steam at 433 K after the CO2 sorption was the only energy source to decompose KHCO3 formed on the AC. The product of the latent heat of steam at 433 K, 2081 kJ/kg H2O, and the accumulated water content of the sorbent during the steam flushing, 85.9 g/kg sorbent (4.77 mol H2O/kg sorbent), yielded the supplied heat source of 178.8 kJ/kg sorbent. The latent heat from steam at 433 K should be allotted to (a) the heat for KHCO3 decomposition, (b) the temperature rise in the sorbent, and (c) the temperature rise in the column vessel. A heat loss through the outer surface of the column was negligible because the column was covered with an annular thermal insulator. The heat of the KHCO3 decomposition at 433 K of 4.02 kJ/kg sorbent could be easily obtained from Figure 6. The mean specific heat of the sorbent between 363 K for the CO2 sorption and 433 K for its release was 1.02 kJ‚kg-1‚K-1 from Figure 7. The mean specific heat multiplied by the temperature difference of 70 K then gave 71.4 kJ/kg sorbent of the heat for the temperature rise of the sorbent. Because the vessel was made of stainless steel (SUS 304), the multiplication of the following parameters gave 105.2 kJ/kg sorbent as the heat for the temperature rise of the vessel, the specific heat of 0.45 kJ‚kg-1‚K-1 for the stainless steel, the vessel weight per unit mass of the sorbent, 3.34 kg/kg sorbent, and the temperature rise of 70 K during the steam flushing step. By adding the heat of the KHCO3 decomposition (4.02 kJ/kg sorbent), the heat for the temperature rise of the sorbent (71.4 kJ/kg sorbent), and the vessel (105.2 kJ/kg sorbent), the total heat required for the CO2 release was found to be 180.6 kJ/kg sorbent. The total heat agreed well with the heat source of 179.8 kJ/kg sorbent, which was supplied by the condensation latent heat of the flushed steam. The heat for the vessel temperature rise dominated the total heat during the CO2 releasing step. Moreover, the substantial heat for the CO2 release, which was obtained to subtract the vessel warming heat from the total heat, mostly consisted of the sorbent warming heat rather than the KHCO3 decomposition heat. From the already mentioned calculation procedure for the heat of the sorbent temperature rise, it should be noted that the lower specific heat, porous materials to support K2CO3 produced the lower heat not only for warming the sorbent, but also for the substantial heat. The hot released gas with CO2 and the excess steam were cooled by a heat exchanger to condense and separate the steam to afford CO2 in high purity, where the heat in the released hot gas was removed by warming a coolant water. From an energyconservative point of view, if the heat held in the hot released gas was recovered by means of any method, the energy consumption for the CO2 recovery could be further reduced. During the sorbent cooling step after the steam flushing step, not only the hot sorbent but also the hot vessel (433 K) should be cooled to below 363 K for the next CO2 sorption. In the same previously described manner, the heat removals from the sorbent and the vessel were estimated to be 71.4 and 126.0 kJ/ kg sorbent, respectively, and the sum of the heat to be removed was 197.4 kJ/kg sorbent. On the other hand, the average temperature of the initial (433 K) and the final (363 K) during the cooling step was 398 K. The product of the evaporation latent heat, 2186 kJ/kg H2O at 398 K, and the removed water
Energy & Fuels, Vol. 20, No. 2, 2006 725 Table 4. Specifications of Commercial-Scale CO2 Recovery Plant item
value
flue gas composition flue gas flow rate flue gas temp & pressure temp for CO2 sorption/release cooling CO2 recovery efficiency CO2 recovery rate concentration of CO2 recovered sorbent and its volume vessel shape
15% CO2, 10% H2O 21 100 Nm3/h 333 K, 0.1 MPa 353 K/423 K (steam) air flow (293 K) 95% 5906 kg/h >99% K2CO3-on-AC, 40.3 m3 × 4 units 3.4 m φ 8 mL, 4 units
amount, 77.0 g H2O/kg sorbent (4.27 mol H2O/kg sorbent) shown in Table 2, then gave 168.3 kJ/kg sorbent for the evaporation heat of the released water. This evaporation heat agreed well with the sum of the heat to be removed of 197.4 kJ/kg sorbent as already mentioned, providing more effective cooling than a forced gas convection without the latent heat transfer. 4.4. Heat Requirement for CO2 Recovery in Practical Plant. Although the heat for the temperature rise of the benchscale vessel was dominant over the total heat for the CO2 release as shown in Table 3, the vessel warming heat in a large-scale plant should be reduced because the vessel weight per unit volume of the sorbent decreased with increasing plant size. As shown in Table 4, the commercial-scale plant with 21 100 Nm3/h of feed gas rate, which was rather smaller than other many plants but large enough for a set of fixed-bed columns, was assumed in order to evaluate the applicability of the K2CO3-on-AC process. To adapt the process to feed gas with a higher rate, the number of fixed-bed column sets should be multiplied as necessary. The other proposed specifications, such as the operating temperature, the CO2 recovery efficiency, and the volume of the sorbent, were presented in Table 4. The flue gas should be compressed to introduce it into the column, and its temperature rises from the initial 333 to 353 K due to the adiabatic compression work. The CO2 sorption/release temperatures were modified to 353/423 K in the commercial plant from 363/433 K in the bench-scale experiments, keeping the same difference between the sorption/release temperatures for both plants. The same temperature difference gives the same CO2 recovery efficiency in the modified temperature range, as demonstrated in the previous result5 for the temperature dependence of the CO2 uptake by the K2CO3-on-AC, while the CO2 recovery energy is conserved by the employment of lower temperature (423 K) steam for the CO2 release in the commercial plant. The heat for the vessel temperature rise in the commercial plant was estimated to be 213 kJ/kg CO2. Dividing the heat of the reaction and that for the temperature rise of the sorbent (Table 3) by the recoverable CO2 amount per unit mass of the sorbent (0.967 mol CO2/kg sorbent in Table 1) could provide the heat per unit mass of the recovered CO2 as 94.1 kJ/kg CO2 and 1680 kJ/kg CO2, respectively. The total heat requirement for the commercial plant, which was composed of the heat of the reaction, the heat for the temperature rise of the sorbent, and that of the vessel, was then expected to be 2033 kJ/kg CO2. Table 5 compares the heat requirement of the CO2 recovery for other processes, in which the K2CO3-on-AC process shows the lowest heat requirement among the amine absorption processes,7-9 demonstrating a remarkable energy reduction for the CO2 recovery with the present process. The Benfield process (7) Erga, O.; Juliussen, O.; Lidal, H. Energy ConVers. Manage. 1995, 36, 387. (8) Kaplan, L. J. Chem. Metall. Eng. 1982, 89, 30.
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Table 5. Comparison of Heat Requirements for CO2 Recovery in Practical Plants process momoethanolamine (conventional) alkanolamine (Gas/Spec FS-1L) hindered amine (Kansai Electric Power Co.) K2CO3 solution (0.14 MPa CO2, Benfield) K2CO3-on-AC (the present work)
heat (kJ/kg CO2) 45457 36938 29309 10301 2033
Table 6. Performances and Cost for Commercial-Scale CO2 Recovery Plant item
value
steam consumption rate electric power cooling water capital cost depreciation rate CO2 recovery cost
6074 kg/ha 870 kWa 418 t/ha 14 million USDb 10%/year 0.102 USD/kg CO2b
a Unit cost: 0.0229 USD/kg steam, 0.0952 USD/kWh, 0.0667 USD/t cooling water. b 1 USD ) 105 Yen.
employing the K2CO3 aqueous solution also shows a lower heat,1 but this heat should not fairly compare with that for the others because the Benfield process only required a higher CO2 partial pressure in the feed gases which led to a lower heat consumption. 4.5. Economic Evaluation of K2CO3-on-AC Process. For the commercial-scale CO2 recovery plant (Table 4), not only the performance, such as the energy consumption rate and the operating energy, but also the capital cost for the plant construction were estimated as shown in Table 6 to estimate the CO2 recovery cost. The steam consumption rate (6074 kg/ h) and the electric power for the flue gas compression blowers and the air-flow cooling pumps, etc. (870 kW), indicated in Table 6 were converted to 12 640 and 3132 MJ/h as a comparable energy consumption rate and to 139.1 and 82.8 USD/h as a cost rate. One should pay attention to the fact that most of the energy consumption and its cost originated in the steam consumption rather than the electric power. The CO2 recovery cost per unit mass of the recovered CO2 was estimated to be 0.10 USD/kg CO2, including the cost for the energy (9) Mimura, T.; Shimojo, S.; Suda, T.; Iijima, M.; Mitsuoka, S. Energy ConVers. Manage. 1995, 36, 397.
consumption (steam, electric power, and cooling water), the redemption of the plant construction, the sorbent cost (0.0142 $/kg CO2), the plant operator’s expense (0.0142 $/kg CO2), and the maintenance expense (0.0058 $/kg CO2). The vacuum cooling step showed a cooling effect slighter than the air-flow cooling in the bench-scale experiments, and therefore this step was omitted in the commercial plant. 5. Conclusion During the CO2 sorption step with K2CO3-on-AC, 54.7% of the stoichiometric amount of CO2 against K2CO3 supported on AC was sorbed, and the corresponding amount of the hydrated water for K2CO3‚1.5H2O was released according to the reaction, K2CO3‚1.5H2O + CO2 ) 2KHCO3 + 0.5H2O. The temperature of the sorbent was raised due to the heat of the reaction for the CO2 sorption. During the steam flushing at 433 K after the CO2 sorption, 92% of the entrapped CO2 was released and followed by moisture separation from the released gas by cooling through a heat exchanger, which gave >99% of pure CO2 gas. The steam supplied not only heat to decompose the KHCO3 formed, but also water to regenerate K2CO3‚1.5H2O by the reaction in reverse. The residual 7% of the entrapped CO2 during the CO2 sorption and the water accumulated on the sorbent were exhausted outside the column under vacuum during the cooling step. The heat required for the CO2 release in the bench-scale column was found to be 4249 kJ/kg CO2, which was composed of the following: the heat for the KHCO3 decomposition, the heat for the temperature rise of the sorbent, and the heat of the column vessel. For expanding the CO2 recovery plant, the heat for the vessel warming per unit mass of the recovered CO2 was minimized and the heat required for the CO2 recovery in the practical plant was estimated to be 1978 kJ/kg CO2. An energysaving effect in the CO2 recovery process from moist flue gases was proposed in comparison with the other absorption methods such as the conventional amine process. The carbon dioxide recovery cost was presumed to be 0.10 $/kg CO2 recovered in the commercial plant. Acknowledgment. Fundamental support from the New Energy and Industrial Technology Development Organization is gratefully acknowledged. EF058027X
