A Test and Demonstration Unit for Concentrating Sulfur Dioxide from

The CCS installation permits smaller flue gas scrubbers to be installed and can potentially save a 400-MW power plant $2.6 million in annual operating...
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Ind. Eng. Chem. Res. 1996, 35, 1409-1416

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A Test and Demonstration Unit for Concentrating Sulfur Dioxide from Flue Gas Joseph P. Dunn, Yeping Cai, Lisa S. Liebmann, and Harvey G. Stenger, Jr.* Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 18015

Dale R. Simpson Department of Earth and Environmental Sciences, Lehigh University, Bethlehem, Pennsylvania 18015

A bench-scale system to continuously separate and concentrate SO2 from flue gas has been developed and tested. The separation is accomplished due to the relative adsorption strengths of SO2 and water on a synthetic mordenite, a phenomenon known as rollup. The continuous concentration system (CCS) is comprised of two pairs of packed beds that cycle between cleaning, roll-up, and regeneration modes. The bed pairs behave identically, with cycles staggered by half of one adsorption-desorption-regeneration period. The cycle for each pair involves regeneration of one bed with dry air, while the other is cleaning the flue gas and subsequently rolling up the SO2. This bench-scale CCS is able to split a 6 standard liters per minute (SLPM) gas stream containing 2000 ppm SO2 into two streams of equal flowrate, one containing 3400 ppm SO2 and the other 70 ppm SO2. The influence of water is strong, with the rollup being optimal at 13% moisture in the flue gas. Adsorption starting with the bed at 75 °C is optimal, and air regeneration at 200 °C and 8 SLPM for 76 min is adequate. An economic analysis of an industrial-scaled CCS capable of treating a 1 000 000 scfm flue gas stream containing 2000 ppm SO2 (typical of a 400-MW coal-fired power plant) was conducted. The CCS installation permits smaller flue gas scrubbers to be installed and can potentially save a 400-MW power plant $2.6 million in annual operating costs. Introduction Sulfur dioxide removal from flue gas is mandated by regulatory and environmental concerns. Removal is also desirable because the presence of sulfur dioxide limits options for further cleanup of the gas such as reduction of nitrogen oxides. However, gas cleanup is hindered by the difficulty in handling large volumes of gas containing a dilute contaminant. Unless a concentration unit for the contaminant is incorporated, the entire gas stream must be treated. This research is directed toward the development of a continuous concentration system (CCS) for SO2 so that a reduced volume of gas can be treated. A reduction of the volume of gas requiring treatment for the removal of SO2 would lead to large capital savings for the SO2 removal apparatus and lower operating costs. It may also permit a smaller retrofit cleanup system where space is severely limited. With SO2 concentrated in a reduced quantity of gas, it enhances the economics of producing sulfuric acid or other sulfur-bearing compounds as salable byproducts. Several ways have been considered to accomplish this concentration goal. These include pressure swing adsorption, temperature swing adsorption [Yang, 1987], and membrane separation. Recent efforts have led to the development of an alternative means to concentrate SO2 into a gas stream of lower flow than the original stream. This technique is best described as chromatographic, sometimes referred to as rollup [Kapoor and Yang, 1987], and utilizes mordenite, a type of molecular sieve, as an adsorbent that concentrates the SO2 by water-driven desorption [Simpson, 1990; and Simpson and Stenger, 1993]. If a gas containing water vapor and SO2, as is typical of flue gas, passes through a bed of mordenite, there is * Author to whom correspondence should be addressed. E-mail address: [email protected]. Phone: (610) 758-5308. Fax: (610) 758-5623.

0888-5885/96/2635-1409$12.00/0

adsorption of SO2 on the mordenite and depletion of SO2 in the effluent stream. As the adsorption bed becomes saturated with SO2, the adsorbed SO2 is spontaneously displaced (rolled up) by water-driven desorption, yielding an exhaust gas enriched in SO2. After rollup, the mordenite is nonsorptive for SO2 until it is thermally regenerated by removal of adsorbed water. Because rollup is a batch phenomenon, it is desirable to develop an operating system for continuous removal of SO2 from flue gas and diversion of a SO2-enriched gas by alternating between several sorptive beds. Experimental Section Figure 1 shows the bench-scale system capable of continuously producing a smaller and more concentrated stream of SO2 from a larger dilute gas stream. The configuration of the beds allows for automatic switching between cleaning, roll-up, and regeneration modes. Both air (house air filtered through a desiccant) and methane (MG Industrial, CP grade) enter a fluid bed of sand held in a furnace (Lindberg) that was operated at 950 °C. This temperature was found in previous studies to result in complete combustion of the methane [Stenger and Meyer, 1992]. The exiting combustion gas is then mixed with 1.5% SO2 in a He stream (MG Industrial, certified) to give a desired SO2 concentration in the flue gas. The methane, air, and SO2 gas flowrates are controlled using three mass flow controllers (Omega, FMA-700 Series). The flue gas is subsequently sent to the CCS where it is divided equally between the two pairs of beds. A metering valve at the exit of each pair of beds allows uniform distribution of the gases between the two pairs, regardless of a possible difference in pressure drop across each bed and its associated tubing. All four sorbent beds are supported vertically and in close proximity. Six three-way and ten two-way solenoid valves (ASCO RedHat) are used to switch flows between beds. All tubing is made of 1/4 in. © 1996 American Chemical Society

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Figure 1. Four-bed continuous concentration system for sulfur dioxide.

Figure 2. Cycle sequence chart for the four-bed CCS. Scheduling for the continuous system was based on the optimal operating conditions for the batch reactor system.

stainless steel and all lines carrying wet flue gas are heated to temperatures higher than the dewpoint of the gas stream. Each of the bed pairs (1-2 and 3-4) operates identically but has cycles that are staggered by half of one “adsorption-rollup-regeneration” period. The cycle for each bed pair involves regeneration of one bed with heated dry air, while the other bed is cleaning the flue gas and subsequently rolling up the SO2. This four-bed design is popular in continuous chromatography, but it is novel for gas separation based on rollup [Stenger et al., 1993b]. The cycle sequence chart is shown in Figure 2. The stainless steel adsorbent bed, illustrated in Figure 3, measures 11 in. length × 3 in. i.d. The distributors are supported by four threaded legs, which can be adjusted to vary the total distributor height and the distance between the mesh screens. Optimal reactor operation was achieved with a 3 in. high distributor having a 2.5 in. separation between the tiers. A 2 in. layer of 0.03 in. diameter glass beads has been placed atop the distributor. The beads improve flow distribution and allow for nearly plug flow operation. Both gas chromatography (GC) and Fourier transform infrared spectroscopy (FTIR) were used to determine the composition of the effluent streams. A GC (Hewlett

Packard Model 5890A) was equipped with a 1/4 in. × 6 ft glass column packed with 80/100 mesh HayeSep R. The separation and quantification were achieved at a temperature of 70 °C with a time interval of 5.5 min. This is accomplished by correctly timing the injections so that the relevant detectable compounds from one injection are eluted from the column during the deadtime of the following injection. All GC sample lines were replaced with nickel tubing and heated with heating tapes to avoid undesired SO2 adsorption and water condensation. A Midac FTIR was also used for SO2 quantification. It has a 10-m-path gas cell (Infrared Analysis, Inc.) with a volume of 3.1 L. At a pressure of 3.0 psia and a temperature of 22 °C no condensation is observed when flue gas is introduced to the cell (see Stenger and Meyer (1992) for FTIR operating details). The computer hardware necessary to perform process control and data acquisition is based on a Zenith Z-Select 100 486SX computer with Kiethley Metrabyte Corp. expansion boards. A DAS-8 analog input board, STA-08 analog input board, and EXP-16 multiplexer/ amplifier are used for reading thermocouple and mass flow input signals. A PIO-12 parallel digital interface board and a SSIO-24 digital I/O board perform solenoid valve positioning and furnace temperature control. A

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Figure 4. Effect of the initial reactor temperature on SO2 breakthrough. Operating conditions: 350 g of mordenite, 3 SLPM, 2000 ppm SO2, 10% H2O. Figure 3. Internal schematic of a packed sorbent reactor. Table 1. Basis Operating Conditions for Four-Bed CCS Experiments -10 + 18 mesh mordenite initial bed temperature feed flowrate feed SO2 concentration feed H2O concentration regeneration temperature regeneration flowrate regeneration time

350 g 75 °C 6 L (STP)/min (continuous) 2000 ppm (2000 × 10-4 %) 13.6% 200 °C 8 L (STP)/min/bed pair 76 min of regeneration + 38 min of cooling

DDA-06 analog/digital expansion board, STA-U output board, and SSIO-24 digital I/O board send analog output signals to mass flow controllers and control the eight cartridge heaters. Labtech Notebook/XE software was used for data acquisition, process monitoring, and realtime control [Dunn, 1995]. In order to facilitate the design of the continuous system, a series of tests were conducted on a single sorbent bed. Design variables tested include the initial reactor temperature, feed moisture content, feed flowrate, and regeneration temperature. The initial reactor temperature was varied between 36 and 75 °C. The feed moisture content, controlled by the air to methane ratio entering the furnace, ranged from 0% to 13.6%. Several flowrates between 2 and 4 SLPM were tested. Sorbent regeneration was conducted at a constant temperature of either 150 or 200 °C. Standard conditions for the single sorbent bed tests are listed in the figure captions. The optimal operating conditions found during parametric studies of the single sorbent bed, listed in Table 1, served as a basis for the operation of the four-bed CCS. With the base operating conditions already defined, the continuous experiments focused on proving the performance and durability of the four-bed continuous system. Results and Discussion Single Sorbent Bed. Effect of the Initial Reactor Temperature. The influence of the initial reactor temperature was investigated by starting operation of

the bed at 36, 46, and 61 °C. The breakthrough curves for SO2 are plotted in Figure 4. It can be seen that a higher initial reactor temperature decreases the mordenite’s capacity for SO2, as evidenced by shorter times until the peak in SO2 concentration. At lower initial reactor temperatures, moderate concentrations of SO2 were found in the leading edge of the roll-up peak. The mordenite surface hydrates quickly at lower temperatures and prematurely forces the SO2 to desorb and break through prior to the rollup. As our process is required to produce clean, SO2-free gas for the period prior to the breakthrough, a higher temperature is preferred. An initial bed temperature of 75 °C (not shown) was deemed optimal for our process. Effect of the Feed Water Concentration. A second important variable is the inlet water concentration, since water concentrations vary significantly depending on the fuel type used during combustion. Figure 5 shows the SO2 breakthrough curves for three different inlet water concentrations: 0%, 8%, and 13.6%. This figure indicates that too little moisture provided an inefficient driving force for the competitive adsorption between SO2 and H2O and thereby decreased the roll-up peak’s height and sharpness. In agreement with a mathematical model of the SO2-H2O-mordenite interaction [Stenger et al., 1993a], there was no SO2 rollup in the absence of water. Increasing moisture decreases the SO2 breakthrough time by lowering the SO2 capacity of the mordenite. This reduced capacity is due to the competitive adsorption of water and SO2 on the mordenite. Effect of the Feed Flowrate. The effect of the total feed flowrate was investigated at 2.4, 3.0, and 4.0 SLPM. As shown in Figure 6, as the flowrate to the system increased, the time until the onset of the rollup decreased. Interestingly, at lower flowrates the roll-up peak became broader and the peak height decreased slightly. This can be attributed to the role of axial diffusion in the sorbent beds. The adsorption capacity of the mordenite for SO2 was similar in all three cases. Effect of the Regeneration Temperature. Regeneration conditions have a dramatic impact on the duration of the clean gas period. As shown in Figure

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Figure 5. Effect of feed moisture content on SO2 breakthrough. Operating conditions: 350 g of mordenite, 3 SLPM, 2000 ppm SO2, initial reactor temperature 75 °C.

Figure 7. Effect of regeneration temperature on SO2 breakthrough. Operating conditions: 350 g of mordenite, 3 SLPM, 2000 ppm SO2, 13.6% H2O, initial reactor temperature 75 °C.

Figure 6. Effect of feed flowrate on SO2 breakthrough. Operating conditions: 350 g of mordenite, 2000 ppm SO2, 13.6% H2O, initial reactor temperature 75 °C.

7, two regeneration temperatures were tested: 150 and 200 °C. Temperatures above 200 °C have been found to irreversibly decrease the adsorption capacity of the mordenite [Stenger et al., 1993a]. At the higher temperature (200 °C) there is a dramatic increase in the capacity of the mordenite, as evidenced by the increase in the time the bed is eluting clean gas. It appears that at the higher temperature adsorbed molecules held tightly in the pores of the mordenite were forced to desorb, thereby exposing more surface sites and increasing capacity. Four-Bed Continuous Concentration System. Clean Gas Stream. Figure 8a is a plot of SO2 concentration versus time for one repeating sequence (see Figure 2: cycle sequence chart) of an experiment using the four-bed CCS. The period between 1254 and 1482 min is characteristic of the performance of the continu-

Figure 8. Experiment with continuous system. (A) SO2 concentration of the clean gas stream. (B) SO2 concentration of the concentrated SO2 stream.

ous system. The concentration profile in Figure 8a repeats every 228 min. The time-average concentration is 70 ppm SO2, and nearly 97% of the feed SO2 has been removed. As Figure 8a shows, the concentration of the clean gas stream never rises over 900 ppm SO2. The peaks in SO2 concentration at 1311, 1425, and 1482 min are due to a slightly premature roll up from the sorbent bed cleaning the flue gas. Concentrated SO2 Stream. The concentration profile for the concentrated SO2 stream is plotted in Figure

Ind. Eng. Chem. Res., Vol. 35, No. 4, 1996 1413 Table 2. Design Basis for Industrial-Scale CCS 1. Power Plant 400-MW coal-fired, burning a 2% sulfur bituminous coal plant is operated 365 days/year flue gas stream 1 000 000 scfm, 200 °F, 1 atm, 2000 ppm SO2, 10% H2O plant

2. Continuous Concentration System three stages, four beds/stage gas streams exiting CCS concentrated 125 000 scfm, 150 °F, 1 atm, 16 000 ppm SO2, 5% H2O clean 875 000 scfm, 150 °F, 1 atm, 0 ppm SO2, 5% H2O mordenite capacity 0.028 g of SO2/g mordenite roll-up ratio 2.0 CCS will cycle between clean gas and concentrated SO2 modes every 1 h 3. Utilities 400 °C and 45 atm (superheated) 199 °C and 15 atm (saturated) cooling water is available from a cooling tower at 30 °C cooling water is returned at 45 °C electricity is purchased off-site steam is available at

Figure 9. Temperature history for a sorbent bed of the CCS.

8b. This figure shows one 228 min repeating sequence, beginning at 1539 min and lasting until 1767 min. The time-average SO2 concentration is 3400 ppm, a 70% increase over the feed SO2 concentration. The depressions in the concentration profile are at the onset and termination of the rollup. The concentration profiles shown in Figure 8 stayed approximately constant throughout the 30 h experiment, with the exception of the start-up period (refer to Figure 2). Several other 24+ h experiments were performed with the continuous system. The effluent concentrations recorded during these experiments were similar to Figure 8. Regeneration Air Stream. Equal flowrates of clean and concentrated gases exiting the continuous system are ensured by a metering valve at the exit of each pair of beds, simplifying the material balance. Comparing the results from the clean gas and concentrated SO2 streams with the known feed concentration, we see that the material balance does not completely close: clean gas (70 ppm) + concentrated SO2 (3400 ppm) 2

*

feed (2000 ppm) This is due to SO2 being desorbed during regeneration. If we assume that all of the SO2 not accounted for in the clean gas and concentrated SO2 streams exits the reactor during regeneration, we can calculate the time-average SO2 concentration of the regeneration air stream:

(2000 ppm - 1735 ppm) × (3 SLPM feed/8 SLPM regeneration) ) 88 ppm SO2 The regeneration air stream has an average SO2 concentration of 88 ppm. This is very similar to the clean gas stream, and it should not present a problem. Temperature History. Sorbent bed temperature can indicate when SO2 breakthrough is about to take place. This may be used in industrial operation to signify when corrections in the timing of sorbent bed switching are necessary. A temperature history for one bed of the continuous concentration system is displayed in Figure 9. The bed temperature is raised from 25 to 210 °C at the start of the regeneration period. For the next 76 min the reactor’s temperature is controlled

around a setpoint of 200 °C. Between 76 and 114 min the cartridge heaters are turned off and the bed cools from 200 to 73 °C by convective heat transfer to the 8 SLPM (23 L/kg of mordenite/min) regeneration air. Flue gas is now sent to the bed, and adsorption begins. The steep temperature rise at the beginning of the adsorption period is due to the highly exothermic adsorption of water. After the adsorption front passes the position of the fixed thermocouple, the temperature drops. At 171 min the effluent gas is switched from the clean gas line to the concentrated SO2 line. The 228 min period was repeated several times. All four beds exhibit similar temperature histories. Economic Feasibility In an effort to lower capital expenditures and operating costs for a 400-MW coal-fired power plant, a preliminary design of an industrial-scale continuous concentration system was developed. Capital and operating costs were calculated for the flue gas treatment section for both a plant with a CCS installed and a plant without a CCS. It was previously found that higher feed concentrations of SO2 did not decrease the roll-up effect [Stenger et al., 1993b]. This important result allows the phenomenon to be cascaded into a multistage process which can increase the concentration of SO2 by severalfold and thus decrease the amount of SO2-containing gas that needs to be treated. The design basis for a three-stage 12-bed CCS is listed in Table 2. For the purposes of the industrial scaleup, the rollup ratio, defined as the average concentration of the concentrated SO2 stream divided by the feed SO2 concentration, has been set at 2.0 instead of 1.7 as found in this work. Previous studies have shown that sorbent beds free of axial dispersion problems will exhibit rollup ratios between 2.0 and 2.4 [Stenger et al., 1993b]. Mathematical models confirm these findings [Stenger et al., 1993a]. The sorbent beds incorporated into the industrial-scale CCS were designed to not show significant axial dispersion, thereby allowing the roll-up ratio of the industrial-scale beds to be conservatively estimated at 2.0. A CCS operating with a roll-up ratio of 2.0 will be able to split a 2000 ppm SO2 stream into two streams of equal flowrate, one SO2 free and the other containing 4000 ppm SO2.

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Table 3. Industrial-Scale CCS Material Balancea

a

Flowrates in cubic ft/min (volume percentage). Note: 0.1 vol % ) 1000 ppm.

Table 4. Total Capital Investment (TCI) for Industrial-Scale CCS (See [Ulrich, 1984] for Description of Method)

Process Description. The operation of the industrial-scale CCS is similar to that of the bench-scale CCS. Flue gas exiting the power plant is sent to the CCS, where it is divided equally between the two pairs of beds. Each of the bed pairs operates identically but has cycles that are staggered by half of one adsorptionrollup-regeneration period. The cycle for each bed pair involves regeneration of one bed with moderate temperature (200 °C) dry air, while the other bed is cleaning the flue gas and subsequently rolling up the SO2. As shown in Table 3, exiting the first stage, the SO2 concentration of the flue gas has doubled, while its flowrate has been halved. Stages two and three operate identically to stage one, with the cycles of the bed pairs being staggered [Stenger et al., 1995]. Exiting stage two, the SO2 concentration is 8000 ppm; exiting stage three, the concentration is 16 000 ppm. Saturated steam at 199 °C is injected between stages one and two

and stages two and three to return the water concentration entering the subsequent stage to 10%. Pressure drop through the shallow sorbent beds is negligible (1 in. of water). Single-stage axial fans are positioned at the entrances to the stack and scrubber to draw the flue gas through the system. Larger backward-curved centrifugal fans are needed for the regeneration air streams. The amount of 30 °C cooling water required by this design is 10.6 m3/min. Superheated steam is supplied to the system at a rate of 982 kg/min. The electric power required by the fans and valve door motors is 828 kW. TCI and TAC of a 12-Bed Continuous Concentration System. The total capital investment (TCI) for the CCS, shown in Table 4, was calculated using standard economic formulas [Ulrich, 1984]. The TCI estimate of $26.1 million is for a complete CCS with all auxiliary facilities installed in a northeastern

Ind. Eng. Chem. Res., Vol. 35, No. 4, 1996 1415 Table 5. Total Annual Operating Cost (TAC) of Industrial-Scale CCS (See [Ulrich, 1984] for Description of Method)

U.S. power plant during the fourth quarter of 1995. As shown in Table 5, of the $11.0 million total annual operating cost (TAC), $4.1 million is assigned to the cost of superheated steam used during regeneration. Scrubber Cost Estimate. Before being released to the atmosphere, the concentrated SO2 flue gas will be treated in a limestone scrubber. The capital and operating cost estimates of the flue gas scrubbers are based on a study conducted by the Electric Power Research Institute [Benitez, 1993]. Capital costs were updated via the CE Index, annual costs were updated via the Consumer Price Index. Plant with a CCS Installation. The total capital cost for a limestone scrubber able to treat a 125 000

scfm gas stream containing 16 000 ppm SO2 (typical exiting conditions for a CCS) is $8.98 million. The yearly operating costs for the scrubber are $3.19 million. These estimates assume a 400-MW plant burning a 2.0% sulfur bituminous coal with 90% sulfur removal. Plant without a CCS Installation. The total capital cost for a limestone scrubber able to treat a 1 000 000 scfm gas stream containing 2000 ppm of SO2 is $47.38 million. The yearly operating costs for the scrubber are $16.85 million. Economic Justification of CCS Installation. To justify the installation of a CCS in the gas treatment section of a power plant, it should be proven that such an installation would lower both capital and

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annual operating costs. Plant without a CCS installed: Capital cost of gas treatment section: Annual operating cost of gas treatment section: Plant with a CCS installed: Capital cost of CCS: Capital cost of scrubber: Total capital cost of gas treatment section: Annual cost of CCS: Annual cost of scrubber: Total annual operating cost of treatment section:

$47.38 million $16.85 million $26.07 million $8.98 million $35.05 million $11.04 million $3.19 million $14.23 million

The CCS has lowered both the capital and operating costs of the gas treatment section of the power plant. Net savings on capital investment allowed by the installation of the CCS are $12.33 million. Net annual operating savings are $2.62 million. Conclusions A four-bed CCS has proven to be capable of using the SO2 roll-up phenomenon to continuously concentrate SO2 in flue gas. The CCS is able to split a 6 SLPM gas stream containing 2000 ppm SO2 into two streams of equal flowrate, one containing 3400 ppm SO2 and the other 70 ppm SO2. In order to make efficient use of the roll-up phenomenon, optimal process operating conditions have been defined. The influence of water is strong with the optimal rollup at 13% moisture in the flue gas. Adsorption starting at 75 °C is optimal. A material balance on the sorbent bed shows that air regeneration at 200 °C and 8 SLPM (23 L/kg of mordenite/min) for 76 min is adequate. Using the results from the parametric studies of the CCS, the optimum set of parameters have been used to determine the capital and operating costs of a system capable of treating 1 000 000 scfm of flue gas at a feed SO2 concentration of 2000 ppm (typical of a 400-MW coal-fired power plant) and producing two effluent gas streams, a 875 000 scfm SO2-free stream and a 125 000 scfm gas stream containing 16 000 ppm SO2. The total capital investment for the industrial-scale CCS is $26.1 million, while the total annual operating costs are $11.0

million. Having less gas requiring treatment allows smaller tail gas scrubbers to be installed. The capital savings gained by the use of the smaller scrubber is $38.4 million. Annual operating savings of $13.7 million are also realized. Net savings on capital investment allowed by the installation of the CCS are $12.3 million. Net annual operating cost savings are $2.6 million. Acknowledgment This work was sponsored by Pennsylvania Energy Development Authority through Contract No. 9203-4013 and Pennsylvania Power and Light Co. The technical and financial support of the sponsors is sincerely appreciated. The authors also thank the Energy Research Center of Lehigh University for partially funding an undergraduate research assistant. Literature Cited Benitez, J. Process Engineering and Design for Air Pollution Control, Prentice Hall, Inc.: Englewood Cliffs, NJ, 1993. Dunn, J. P. A Test and Demonstration Unit for Concentrating SO2 from Flue Gas. M.S. Thesis, Lehigh University, Bethlehem, PA, 1995. Kapoor, A.; Yang, R. T. Roll-up in Fixed-Bed, Multicomponent Adsorption Under Pore-Diffusion Limitation. AIChE J. 1987, 33 (7), 1215. Simpson, D. Mordenite and Mordenite Aggregate Synthesis. U.S. Patent 4,935,217, 1990. Simpson, D.; Stenger, H. G. A Process for Removing Sulfur Oxides from a Gas Stream. U.S. Patent 5,223,237, 1993. Stenger, H. G.; Meyer, E. C. Laboratory-Scale Coal Combustor for Flue Gas Emission Studies. Energy Fuels 1992, 6, 277. Stenger, H. G.; Simpson, D. R.; Hu, K. Competitive Adsorption of NO, SO2 and H2O onto Mordenite Synthesized from Perlite. Gas Sep. Pur. 1993a, 7 (1), 19. Stenger, H. G.; Hu, K.; Simpson, D. Chromatographic Separation and Concentration of Sulfur Dioxide in Flue Gases. Ind. Eng. Chem. Res. 1993b, 32, 2736. Stenger, H. G.; Simpson, D. R.; Cai, Y.; Dunn, J. P.; Liebmann, L. S. A Final Report to PEDA: A Test and Demonstration Unit for Concentrating SO2 from Flue Gas; Pennsylvania Energy Development Authority: Harrisburg, PA, 1995. Ulrich, G. D. A Guide to Chemical Engineering Process Design and Economics; John Wiley & Sons, Inc.: New York, NY, 1984. Yang, R. T. Gas Separation by Adsorption Processes; Butterworth Publishers: Stoneham, MA, 1987.

Received for review September 21, 1995 Accepted January 4, 1996X IE950585J X Abstract published in Advance ACS Abstracts, February 15, 1996.