Carbon Filter Process for Flue-Gas Carbon Capture on Carbonaceous

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Carbon Filter Process for Flue-Gas Carbon Capture on Carbonaceous Sorbents: Field Tests of Steam-Aided Vacuum Swing Adsorption Bryce Dutcher, Kaspars Krutkramelis, Hertanto Adidharma,* and Maciej Radosz Soft Materials Laboratory, Department of Chemical and Petroleum Engineering, University of Wyoming, Laramie, Wyoming 82071, United States ABSTRACT: A carbon filter that selectively captures CO2 from flue gas on porous carbonaceous sorbents is demonstrated in field tests. A new sorbent regeneration process, referred to as steam-aided vacuum swing adsorption (SA-VSA), uses steam under vacuum to displace CO2 from the carbon. Steam aids vacuum by reducing the partial pressure of CO2 without the need to reduce the total pressure to deep vacuum, which is a challenge for vacuum swing adsorption (VSA). In turn, vacuum prevents the bulk condensation of steam on the sorbent, which is a challenge for CO2 recovery using direct steam. Over 100 sorption−desorption cycles on flue gas produced in the lab and at two coal-fired power plants demonstrate that the SA-VSA-equipped carbon filter process can produce a nearly pure CO2 product (at least 98%) while achieving high recovery (at least 98%). While stable, flexible, and robust in achieving the very high recovery and purity targets, this technology offers ample room for improvement through process optimization and, especially, sorbent optimization. For one of the reasonable but arbitrarily selected and inexpensive sorbents, a preliminary cost estimate example suggests an energy penalty of 31%, which should be less or much less for better sorbents.

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INTRODUCTION CO2 separation from combustion flue gas is by far the most expensive step in the carbon capture and sequestration (CCS) process. Common separation approaches that are technically feasible, such as amine or ammonia absorption, pressure swing adsorption, and membrane separation, tend to be expensive because of high regeneration costs or the need to compress or refrigerate the flue gas.1−4 A low-pressure adsorption on porous carbon, referred to as a carbon filter process,5 on the other hand, requires neither compression nor refrigeration. Porous carbon sorbents are selective for CO2 relative to nitrogen, a major inert flue gas component. Studies suggest that adsorption using activated carbons has lower regeneration energy than other physical sorbents.6 Therefore, a carbon filter process has potential to be more cost-effective. In the authors’ previous work,7 several methods of regenerating an activated carbon bed were examined. In one method, using direct steam for temperature swing adsorption (TSA), a rapid CO2 desorption led to two fractions: a CO2-lean fraction, similar in composition to flue gas, and a CO2-rich fraction that was nearly 100% pure, which is an advantage. However, in this approach, condensed steam affected the sorbent capacity for CO2, not to mention that the hot bed had to be cooled before the next sorption cycle, which can be slow. In another regeneration method, namely, vacuum swing adsorption (VSA), CO2 was removed under vacuum, without the presence of water, which led to similar two fractions: CO2lean and CO2-rich fractions. However, to desorb the sorbent completely, pressures much below 20 Torr were necessary, which is impractical for a large-scale operation. To alleviate the issues associated with TSA and VSA, in the same previous work, the authors proposed a hybrid approach referred to as steam-aided VSA (SA-VSA). In SA-VSA, moderate vacuum alone removes an initial fraction of sorbate, and more © 2012 American Chemical Society

important, if properly chosen, it can essentially prevent bulk condensation of the steam. Steam, on the other hand, serves as a displacement medium for the sorbed CO2 and, hence, eliminates the need to reach very low pressures. In addition to efficient sorbent regeneration, with no need for extremely low pressures and no significant steam condensation, SA-VSA demonstrated a potential for nearly complete recovery of nearly pure CO2 while allowing for flexibility in optimizing CO2 recovery and purity by adjusting the two-fraction desorption cutoff. That previous work on SA-VSA provided a crucial proof of concept on a synthetic CO2−nitrogen mixture, but it did not demonstrate how a common sorbent can perform on a real flue-gas mixture in a series of cycles, including recycle of the CO2-lean fraction toward a complete recovery of nearly pure CO2. Therefore, the objectives of this work are (1) to build a trailer-mounted field test unit that can be transported to power plants, (2) to demonstrate that SA-VSA can recover at least 95% of CO2 from combustion flue gas and that the purity of recovered CO2 will be at least 95%, which requires recycling the CO2-lean fraction, (3) to explore the impact of water accumulation in multiple adsorption−desorption cycles, and (4) to generate a back-of-the-envelope example of relative costs for the adsorbents used. This unit should be fully automatic to test the sorbent stability in at least 100 consecutive cycles. To provide accurate data for the material balance of the major components and avoid significant sampling perturbations associated with water vapor in combustion flue gas, water vapor will be condensed and separated from the flue gas prior to charging the sorbent tanks, which is consistent with the Received: January 29, 2012 Revised: March 15, 2012 Published: March 20, 2012 2539

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the vessel on the right, desorption occurs using steam under vacuum to displace CO2 from the sorbent. The initial, CO2-lean fraction is sent back into the feed through a recycle tank. The final, CO2-rich fraction is captured as product. The field test unit is shown in Figure 2. A detailed flow diagram, including all of the major equipment components, is shown in Figure

process concept. This work is not aimed at sorbent selection, process optimization, SOx and NOx impact, and extensive economic evaluation.

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MATERIALS AND METHODS

Initial lab-demonstration tests of the field test unit use synthetic flue gas, 12% CO2 with the balance of air. Pure CO2 is purchased from United States Welding, Inc. When tested in the field, the unit uses combustion flue gas taken from a slip stream in the power plant stack. Field tests are performed using a coal-derived sorbent, AC1. AC1 is composed of spherical particles, 1.5 mm in diameter. The bulk density is 520 kg/m3, and the void fraction is 0.36. AC1 is chosen because of its sorption properties, primarily sorption capacity. Preliminary economic analyses are performed using AC1 and another coal-derived sorbent, ACC. ACC is composed of cylindrical particles, 3.5 mm in diameter and on average 7 mm in length. Its bulk density is 490 kg/m3, with a void fraction of 0.56. Sorption data for both sorbents are

Table 1. Sorption Properties of ACC and AC1 at Room Temperature sorbent

ACC

AC1

CO2 sorption capacity at 1 bar (wt %) CO2 sorption capacity at 10 bar (wt %) N2 sorption capacity at 1 bar (wt %) N2 sorption capacity at 10 bar (wt %) selectivity at 1 bar selectivity at 10 bar BET surface area (m2/g) micropore surface area (m2/g) external surface area (m2/g) micropore volume (cm3/g) mesopore volume (cm3/g) total pore volume (cm3/g)

6.7 18.6 0.5 2.2 13.4 8.5 1107 545 562 0.246 0.227 0.473

9.9 26.0 1.0 5.3 9.9 4.9 821 602 219 0.280 0.171 0.451

Figure 2. Field test unit for SA-VSA. 3. Each of the two sorbent vessels made of 1.5 mm thick stainless steel is 4.6 L in volume. They are 14 cm in diameter and 34 cm in height. The ends are nearly hemispherical. Each vessel is covered with 3 cm thick foam rubber insulation to minimize heat losses. The vacuum pump is a KNF N820.3FTP diaphragm pump, rated for 20 L/min at atmospheric conditions. The flue gas preparation section consists of a dryer that condenses the bulk of water, a guard bed that captures minor impurities, and a compressor that pulls a particulate-free slip stream of flue gas from the power plant stack to a feed tank. The purpose for the compressor/tank assembly is to accomplish a stable flow control toward an accurate material balance in the field test. In the commercial process, a simple blower is envisioned rather than a compressor and feed tank. The dryer is used to prevent water damage to downstream equipment, including the guard bed, mass flow controllers, and gas analyzers. Water vapor removal from the flue gas is also to simplify the water mass balance assessment during the operation cycle and to enable us to investigate accurately the impact of water accumulation in multiple adsorption−desorption cycles. The guard bed removes trace components, such as NOx, SO2, CO, and mercury. The guard bed, containing sacrificial activated carbon saturated with CO2 at the operating conditions, is overdesigned to ensure a long life over many cycles. The breakthrough of trace gases, if any, is continuously monitored at the inlet to the field test unit. An Agilent 5975C mass spectrometer, configured as a real time gas analyzer (RTGA) by Diablo Analytical, Inc., monitors the desorption stream composition. The RTGA is calibrated using 50, 20, and 12% CO2 (balance N2) certified gas mixtures from Airgas, as well as pure CO2 and N2. Compositions of all other streams are measured with Alphasense sensors, using chemical sensors for O2 and trace gases and infrared sensors for CO2; the N2 concentration is taken as the balance. All equipment components are interfaced with the main computer using National Instruments DAQ board and LabVIEW software. This computer collects composition, flow rate, temperature, and pressure data. The same computer controls valves and flow controllers. In brief, the entire unit is fully automated for a continuous operation from startup to shutdown. During the adsorption stage, flue gas at approximately 32 °C and 740 Torr, both slightly higher than ambient conditions, flowing at a rate of 2 L standard temperature and pressure (STP)/min (in this

provided in Table 1. Sorption capacity under flue gas conditions varies significantly depending upon process parameters, as discussed in our previous work.7 Experiment. A simplified flow diagram of the field test unit that illustrates the process concept is shown in Figure 1. Two sorbent beds alternate between sorption and desorption stages toward a continuous process. Flue gas enters one vessel, in this case, the vessel on the left, where CO2 is removed, and the cleaned flue gas is sent to the stack. In

Figure 1. Simplified flow diagram for the field test unit. 2540

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Figure 3. Flow diagram of the automated pilot unit: CO is a carbon monoxide sensor; CO2 is a carbon dioxide sensor; FC is a flow controller; FM is a flow meter; NO is a nitrogen monoxide sensor; NO2 is a nitrogen dioxide sensor; O2 is an oxygen sensor; P is a pressure bypass valve; PT is a pressure transducer; R is a water reservoir; SO2 is a sulfur dioxide sensor; TC is a thermocouple; and WM is a water meter. work, STP is defined as 25 °C and 1 atm), is passed through the sorbent bed. This flow rate is chosen to synchronize adsorption with desorption. The bed temperature is approximately 33 °C, which remains constant naturally. The flue gas composition is monitored to make sure that it also remains nearly constant for a given power plant (but it may differ from plant to plant). The effluent gas is cooled to recover water for material balance calculations and then vented. The beds are switched when CO2 breaks through, which, for these tests, is declared when the CO2 concentration in the effluent reaches 2% (breakpoint). The desorption stage begins by pulling vacuum. When the vessel reaches the desired pressure, selected to be 30 Torr based on our previous work,7 the pressure control valve admits steam to maintain a constant pressure (of 30 Torr). The gas discharged from the vacuum pump is dried in a condenser for a material balance. The composition of this stream is monitored with the RTGA. The initial CO2-lean fraction, rich in N2 and O2, accumulates in the recycle tank and eventually blends with the feed stream. When the cutoff is adjusted between the recycle and product fractions, it is possible to control the purity of recovered CO2. After the effluent stream has been monitored for material balance, it is vented. To provide steam, a peristaltic pump feeds water from a reservoir into an oven. The steam leaves the oven at 150 °C, and it becomes saturated when it reaches the pressure control valve. The pressure control valve allows for a small amount of steam to enter the desorption vessel when the vessel pressure reaches the target vacuum. The valve is heated to 130 °C, which ensures no liquid is being pulled into the system. The excess steam is condensed and returned to the reservoir. The pressure control valve doubles as a flow meter to measure the amount of water entering the system. Water exiting the condensers for the adsorption effluent and pump discharge drains into a float trap. The water flow rate out of the trap is measured.

In separate experiments, evacuation curves, which are used to obtain certain sorbent properties needed to estimate relative costs for the two adsorbents studied, are measured for five vessels as follows: 7 L volume with 7.6 cm diameter, 4.6 L volume with 14 cm diameter, 0.62 L volume with 7 cm diameter, 0.5 L volume with 4 cm diameter, and 0.15 L volume with 4 cm diameter. For each of these vessels, three evacuation curves are taken, using the empty vessel, the vessel filled with clean sorbent, and the vessel filled with sorbent saturated with 12% CO2 in N2 at ambient conditions. All evacuation curves are taken at ambient temperature. To obtain the needed sorbent properties from these evacuation curves, preliminary approximations of the evacuation processes are used, which are described below. Preliminary Approximations. In the spirit of a first-pass, back-ofthe-envelope approximation, models for evacuation processes of an empty vessel, a vessel filled with clean sorbent, and a vessel filled with sorbent saturated with CO2 are developed. The rate of pressure change for an isothermal empty vessel evacuation is derived from the ideal gas mole balance dP nRT ̇ =− dt V

(1)

where P is the vessel pressure, ṅ is the gas molar flow rate of the vacuum pump, which is a function of the pressure derived from data provided by the manufacturer, R is the molar gas constant, T is the absolute temperature, V is the vessel volume, and t is the time. For a vessel filled with clean sorbent, the gas molar flow rate reduces considerably because of resistance to flow through the sorbent. A simple model can be applied by introducing a correction factor r (r > 1), which reflects the magnitude of the resistance to flow, such that the gas molar flow rate of the vacuum pump is now given by ṅ/r. When the pressure of the system is assumed uniform at any time and mole 2541

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balance is applied about the tank, the rate of pressure change for an isothermal evacuation of a sorbent-filled vessel is given by dP nRT ̇ 1 =− dt Vε r

(2)

where ε is the void fraction of the sorbent. The dimensionless resistance r is fitted to the experimental evacuation data. For an isothermal evacuation of a vessel filled with sorbent saturated with 12% CO2 in N2, the mass transfer of gases from the sorbent complicates the process further. To simplify this problem, to a first approximation, we neglect the mass transfer of N2 and write the total mole balance of gas in the tank

⎡K a dP ṅ ⎤ = ⎢ c (C* − C) − ⎥RT ⎣ Mε dt V εr ⎦

(3) Figure 4. Typical breakthrough curve of CO2 for the field test unit.

where Kc is the overall mass-transfer coefficient, a is the interfacial area per unit volume of the bed, M is the molecular weight of CO2, C is the CO2 concentration in the bulk gas phase, and C* is the concentration of CO2 in equilibrium with the sorbent loading q; C* and q are related by the isotherm. In our case, because N2 desorbs quickly at the beginning of the process and, thus, most of the time during the evacuation process, only CO2 desorbs from the sorbent, the assumption of no mass transfer of N2 for the entire process should be reasonable. Equation 3 is solved simultaneously with the mass balance equations for CO2 in the fluid and solid sides

Ka dC nRT ̇ = c (C* − C) − C dt ε PV εr ρ(1 − ε)

dq = − K ca(C* − C) dt

(4) (5)

where ρ is the sorbent particle density. Kca is fitted to the experimental evacuation data. Note that, in these preliminary approximations, uniform pressure and bulk gas composition at any time have also been assumed. The error as a result of these assumptions is expected to be insignificant because the pressure drop across the bed is very small and, most of the time during the evacuation process, the mole fraction of CO2 is almost 1.

Figure 5. Recovery for consecutive cycles for the Pawnee Station tests.

Figure 6, which shows a typical concentration of CO2 in the desorption stream. To achieve this purity, for example, the

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RESULTS In this work, CO2 recovery is defined as the volume of CO2 in the product divided by the volume of CO2 fed to the vessel and CO2 purity is defined as the volume fraction of CO2 in the product. A total of 110 cycles are performed with the field test unit at three different locations, on the same sorbent with no cleanup between the locations. The first 30 cycles are performed in the laboratory using a synthetic flue gas of 12% CO2 with the balance air. The next 50 cycles are performed at Jim Bridger Power Plant, owned by PacifiCorp Energy, in Rock Springs, WY. The dry flue gas at Jim Bridger is 11% CO2, 11% O2, with the balance N2. The final 30 cycles are performed at Pawnee Station, owned by Xcel Energy, in Brush, CO. Pawnee Station’s dry flue gas is 14% CO2, 7% O2, with the balance N2. High CO2 recovery is observed in all cycles. The average recovery is 98.4%. This is because the only time CO2 is lost from the system is in the adsorption effluent just before the breakpoint. This is illustrated in Figure 4. Recovery can be adjusted easily by changing the breakpoint. To achieve this recovery, the breakpoint is set at 2% CO2 in the effluent. Because the CO2 recovery can be adjusted in this manner, it is not a function of the sorbent CO2 capacity. Recovery is shown for the Pawnee Station tests in Figure 5. The CO2 purity of each cycle is also demonstrated to be very high, on average 98.2%, and easily adjusted by changing the cutoff point between recycle and product. This is illustrated in

Figure 6. Typical concentration of CO2 in the desorption stream of a single cycle.

cutoff is set when the concentration of CO2 in the desorption stream reaches 75%. The recycled fraction is approximately one-third of the total desorbed gas by volume. The average CO2 concentration of the recycle stream is always around 35%. Similar to recovery, purity therefore is not a function of the sorbent capacity for CO2 either. Figure 7 shows purity for consecutive cycles for the Pawnee Station tests. Figure 8 shows the concentration of major components in the desorption stream over several consecutive cycles, illustrating the stability of SA-VSA. 2542

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Figure 7. Purity for consecutive cycles of the Pawnee Station tests.

Figure 9. Mass of water accumulated in each cycle for the Pawnee Station tests.

Figure 8. Concentration of major components in desorption stream over several cycles. Results are for the Pawnee Station test. Figure 10. Time required for each cycle for the Pawnee Station tests.

In previous work,7 it was observed that, for 15 cycles, some steam used for desorption was retained in the bed from cycle to cycle. It was hypothesized that this accumulation was due to either slight adsorption on the sorbent or trace condensation. If the former was true, the sorbent would eventually become saturated and the accumulation would cease. If the latter was true, preventing heat loss would reduce the accumulation, which, indeed, turned out to be the case. After 110 cycles in this work, the trace condensation on average is found to be about 4.8 g/cycle, which corresponds to 0.08% of the sorbent weight. This is illustrated for the Pawnee Station cycles in Figure 9. As it turns out, such low water accumulation has only a slight effect on the CO2 capacity of the sorbent, observed as a small but steady decrease in cycle time, as shown in Figure 10. Because recovery and purity are not functions of the CO2 capacity, however, they are not affected by this level of water accumulation. Because the time required for adsorption is easily adjusted by changing the feed rate of flue gas, the cycle time is limited by the desorption rate. For AC1 in the vessels used for the field tests, the cycle times typically range from 35 to 50 min, decreasing slightly over the course of 110 cycles. To debottleneck the process, that is, to reduce the desorption time, a better understanding of the desorption rate and better sorbents will be needed, which is the subject of the next section. Resistance to Flow and Mass Transfer. Equation 1 is used with the experimental data from the empty vessels to verify ṅ derived from the data of the vacuum pump

manufacturer. We find that the model with ṅ obtained can well-represent the evacuation process. When evacuation curves of empty vessels are compared to those of filled vessels, it is clear that the flow rate of the pump is affected by a resistance to flow because of the sorbent. This dimensionless resistance r is determined for two reasonable but arbitrary sorbents, AC1 and ACC, by fitting the theoretical evacuation curve of a vessel filled with clean sorbent to the experimentally measured pressure curve. The resistance, which is constant for a given sorbent, is 15.4 for AC1 and 6.8 for ACC. This resistance indicates that the time and, thus, energy required to decrease the pressure in a bed of ACC is less than half that of an equal volume of AC1. While AC1 has a somewhat higher capacity for CO2, ACC has a much lower resistance to flow during the desorption stage. This means that the operating costs for ACC will likely be much lower than those for AC1. Another important parameter for scale-up is the masstransfer coefficient. In general, this coefficient is a function of the particle geometry and properties, the flow condition, and the fluid properties at the process condition. It is observed that the mass-transfer coefficient for a particular sorbent reaches an asymptotic value and does not change significantly for vessels of large volume, such as 4 L or more, regardless of vessel dimensions, as shown in Figure 11. This is expected because the flow condition will not change much as the volume of the vessel becomes large. For the purposes of scale-up, in this preliminary estimation, the asymptotic value is used as the 2543

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and minor components, such as fans. In this estimation, we assume that the cost of both sorbents is the same. In the first scale-up case, used to determine the operating costs, a CO2 capture rate is assumed to be 1000 tons/year. If AC1 is used, the best operating case results in a vessel that is 3.1 m in diameter and 4.5 m in height. A vacuum pump capacity of 1879 L/min at standard conditions is required. The cycle time is 180 min. The total operating cost is $55/ton of CO2. A total of 89% of this cost is associated with the vacuum generation. However, if ACC is used, the vessel size becomes 3 m in diameter and 2.2 m in height. With the same vacuum pump capacity, the cycle time is reduced to 80 min, and hence, the total operating cost becomes $21/ton of CO2, of which 78% is associated with the vacuum generation. For the 15 year cash flow, we assume a 15 year project life with a 2 year construction period. MACRS108 depreciation is used with a 35% tax rate. The interest rate is assumed to be 15%. The cost of recovered CO2 in this scenario is estimated assuming that the net present value of the project is zero, which implies, somewhat unrealistically, that the CO2 market value is the same as the cost of its recovery. The case study is performed for the Wyodak Power Plant near Gillette, WY, which was chosen for previous work.9 It is a 335 MW coal-fired power plant, which produces 3 520 000 tons of CO2/year. An arbitrary goal of this case study is to capture 90% of CO2. To accomplish this with AC1, a capital investment of $558 million is required. In combination with the operating cost for a 15 year cash-flow period, the cost of recovered CO2 is $81/ton of CO2. The energy penalty for this process, that is, the amount of energy used relative to the amount of energy currently produced by the plant, is 75%. If this process is performed with ACC, although the capital cost is only $349 million, the cost of recovered CO2 is $47/ton of CO2 and the energy penalty is only 31%. This simple example suggests that, under special circumstances, a lower capacity sorbent may result in a more economic CO2 capture, even if it costs the same. However, a lower capacity sorbent can, in fact, be much less expensive, which would further improve its economic attractiveness. More important, this work suggests that there is a substantial room for reducing the costs through process optimization and, especially, sorbent optimization. We envision an ideal sorbent to have low resistance to flow, low sorption capacity for water, large CO2 working capacity in the presence of water, and tolerance for temperature fluctuations.

Figure 11. Effect of the vessel volume on the mass transfer (Kca) of AC1.

mass-transfer coefficient and taken to be a constant for each sorbent. The value of Kca is 0.026 min−1 for AC1 and 0.146 min−1 for ACC. Figure 12 shows that our first-pass approach,

Figure 12. Desorption model fitted to experimental data for evacuation of a 4.6 L vessel filled with saturated sorbent.

however approximate, adequately represents the experimental data for the evacuation of a vessel filled with saturated sorbent. This confirms that the assumption of uniform pressure and gas composition at any time is reasonable. Because the difference between the ambient temperature and the operating temperature is small, there is no need to correct the mass-transfer coefficient obtained at ambient conditions because it will cause negligible error in the preliminary cost estimates. Operating Costs and Energy Penalty. Even though this work is not aimed at accurately estimating the cost of carbon capture and comparing it to other separation methods, we suggest an approximate example of relative costs for the two sorbents used. Two cases are considered: a case study using only the operating cost, solely based on utility consumption, and a case study using 15 year cash flow, which accounts for capital cost as well as the operating cost. The operating cost is calculated by estimating the consumption of electricity for the vacuum pump and fans, steam for sorbent regeneration, and cooling water for steam condensation. Prices are assumed to be $0.07/kWh for electricity, $0.43/1000 lbs for steam independent of the main steam cycle of the power plant, and $0.01/1000 gallon of cooling water. The capital cost is based on the vessel sizes discussed below, $1650/ton of sorbent, vacuum pumps,

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CONCLUSION

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AUTHOR INFORMATION

SA-VSA is shown to be an efficient means of capturing CO2 from flue gas. Moderate vacuum prevents bulk steam condensation, while steam sweeps CO2 and, hence, makes it unnecessary to reach deep vacuum as in a typical VSA. Field tests demonstrate that the SA-VSA version of the carbon filter process can recover over 98% of nearly pure CO2 from flue gas in a stable operation over 100 cycles. This work also demonstrates that the cost of recovered CO2 hinges on the ease of desorption. For a reasonable but somewhat arbitrary sorbent, the energy penalty is found to be as low as 31%, but process and sorbent optimization can reduce it substantially.

Corresponding Author

*Telephone: 307-766-2500. Fax: 307-766-6777. E-mail: [email protected]. 2544

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Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work was funded by Wyoming’s Enhanced Oil Recovery Institute, Supercritical Fluids LLC, the state of Wyoming’s Clean Coal Program administered by the University of Wyoming’s School of Energy Resources, the Electric Power Research Institute, Pacificorp Energy, Xcel Energy, and a discretionary fund of one of the authors (Maciej Radosz). The authors also thank Mr. Ryan Taucher, Pacificorp Energy’s Jim Bridger Power Plant, WY, Mr. Barry Andrews, Xcel Energy’s Pawnee Station, CO, and Dr. Xin Hu, who characterized the sorbents.

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NOMENCLATURE a = interfacial area per volume of bed (m2/m3) C = concentration (kg/m3) C* = concentration in equilibrium with sorbent loading (kg/ m3) ε = void fraction of bed Kc = overall mass-transfer coefficient (m/s) M = molecular weight (g/mol) ṅ = molar flow rate (mol/s) P = pressure (Pa) q = sorbent loading (g of sorbate/g of sorbent) R = molar gas constant r = resistance to flow ρ = apparent density of sorbent (kg/m3) T = temperature (K) t = time (s) V = volume (m3)

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REFERENCES

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