Combustion of Low-Concentration Coal Bed Methane in a Fluidized Bed

Feb 20, 2011 - ... Shapingba District, Chongqing 400030, People's Republic of China. § ... of British Columbia, Vancouver, British Columbia V6T 1Z3, ...
2 downloads 0 Views 1MB Size
Download PDF
ARTICLE pubs.acs.org/EF

Combustion of Low-Concentration Coal Bed Methane in a Fluidized Bed Zhongqing Yang,†,‡ John R. Grace,§ C. Jim Lim,§ and Li Zhang*,†,‡ †

Key Laboratory of Low-Grade Energy Utilization Technologies and Systems, Ministry of Education, and ‡College of Power Engineering, Chongqing University, Shapingba District, Chongqing 400030, People’s Republic of China § Department of Chemical and Biological Engineering, University of British Columbia, Vancouver, British Columbia V6T 1Z3, Canada ABSTRACT: Low-concentration coal bed methane combustion tests were carried out in a bubbling fluidized-bed reactor with a mixture of limestone and crushed catalyst as bed material. The effects of bed temperature (723-923 K), inlet methane concentration (0.15-3 vol %), and superficial velocity (0.1-0.25 m/s) on methane conversion were studied. Gas was sampled from the bed and the freeboard to analyze the axial profiles of methane in the reactor. With kinetic parameters obtained in a fixed-bed microreactor, predictions from a two-phase model were compared to the experimental data from the fluidized bed. The results showed that bed temperature greatly affected the conversion, with a higher methane conversion obtained by increasing the temperature, reducing the inlet methane concentration, and decreasing the superficial gas velocity.

1. INTRODUCTION Coal bed methane (CBM) is a form of natural gas extracted from coal beds during mining. Low-concentration CBM, in which the methane volumetric concentration is less than 5%, is difficult to capture because the air volume is large and the methane resource is dilute and variable in concentration and flow rate. The explosive range of methane is 5-15% by volume for atmospheric air. For mining safety, the CBM is always diluted to less than 5% and even lower than 1% in ventilation air. At present, almost all low-concentration CBM in the world is emitted to the atmosphere without treatment.1 In China, the total amount of methane emitted to the atmosphere is 20  109 m3 every year, a value which is increasing with the growth in coal production.2 These emissions cause energy waste of a nonrenewable resource, as well as intensified greenhouse gas emission, given that methane has a greenhouse gas potential of 21-23 times that of CO2. Therefore, it is essential to find a way to use low-concentration CBM. Catalytic combustion of methane is a clean and efficient way of burning methane compared to conventional flame combustion. Full oxidation of methane can be achieved on suitable catalysts at low temperatures. Moreover, catalytic combustion enables air/ fuel mixtures to be burned well outside the flammability limits.3 Therefore, it provides a very attractive way to recover energy from low-concentration CBM. Kinetic mechanisms of methane catalytic combustion involve multi-step surface reactions.4 However, the overall combustion methane reaction can be represented simply by

temperature. The reaction order for methane was found to be between 0.5 and 1. Chaouki et al.7 studied a cyclic fixed-bed reactor with reversed flow using Pd/Al2O3 (0.2 wt % palladium) as a catalyst and found that an appropriate choice of operating parameters allowed for lean methane combustion at a temperature of 923 K with no NOx production. Gosiewski and Warmuzinski8,9 employed the flowreversal principle to burn mine ventilation air. They found that the rise in bed temperature usually led to the formation of hot spots and deactivation of the catalyst, even causing homogeneous combustion. One of the most important reasons for catalyst deactivation is heat, which cannot be withdrawn effectively from the catalyst, causing overheating and thermal deactivation.10 It has been suggested11 that heat extraction from the catalytic combustion system could be conveniently accomplished in a fluidized-bed reactor. In addition, fluidized-bed reactors can be operated isothermally without developing hot spots.12 Zukowski13 investigated gaseous fuel catalytic combustion in a fluidized-bed reactor, operating in the bubbling regime, with manganese oxides as the bed material. With bed temperature > 1023 K, the process consisted of heterogeneous reactions on the surface of the catalyst particles and homogeneous reactions in the interstices between particles. Hayhurst et al.14 studied propane combustion on Pt pellets and compared this to sand. Sand was found to inhibit the combustion, whereas platinized alumina particles catalyzed combustion. Foka et al.15 and SotudehGharebagh16 investigated catalytic premixed combustion of natural gas, with an inlet mixture containing 4% methane at temperatures of 673-873 K. They found that the turbulent flow regime was most suitable for combustion of natural gas. Iamarino10 employed a Cu/γ-Al2O3 catalyst to study a premixed fluidized-bed reactor under lean conditions. They reported that

CH4 þ 2O2 f CO2 þ 2H2 O 0 ¼ - 802 kJ=mol ð1Þ ΔH298 Platinum and palladium are generally accepted as the most active catalysts for low-temperature combustion.5 Lee and Trimm6 summarized activation energies and reaction orders from various experiments. It was noted that activation energies were quite variable, being dependent upon the catalyst and r 2011 American Chemical Society

Received: November 21, 2010 Revised: January 16, 2011 Published: February 20, 2011 975

dx.doi.org/10.1021/ef101573y | Energy Fuels 2011, 25, 975–980

Energy & Fuels

ARTICLE

methane could be completely converted at temperatures below 973 K in a bubbling fluidized bed of 1 mm catalyst particles. Few authors have reported the influence of operating conditions on lean methane combustion in fluidized-bed reactors. For low-concentration CBM, the flow and methane concentration vary widely according to the location and technique of mining. In the present work, low-concentration methane catalytic combustion was carried out in a bubbling fluidized-bed reactor, with bed temperature, inlet methane concentration, and superficial gas velocity as variables. Methane profiles were measured to obtain a better understanding of the combustion. In addition, a two-phase model, combined with kinetic data, is presented to describe the reactor behavior.

2. EXPERIMENTAL SECTION Because palladium (Pd) is one of the most active catalysts for low temperature, a commercial G-74D catalyst (S€ud Chemie, Inc.) with its main active component 0.5 wt % Pd on an alumina carrier was used for the experiments. The surface area of the catalyst is 210.7 m2/g, and the pore volume equals 0.437 cm3/g. The low-concentration CBM combustion tests were carried out in a bubbling fluidized-bed reactor with limestone and catalyst as bed materials. To assess the catalytic activity, kinetic tests were also conducted in a fixed-bed microreactor. 2.1. Kinetic Tests. The kinetic study of methane oxidation on the catalyst was carried out in a laboratory fixed-bed microreactor of 9.5 mm outer diameter and 95 mm length. The G-74D catalyst was crushed and screened to 180-300 μm. A total of 1.5 mg of the catalyst was mixed with γ-Al2O3 powder (80 μm average particle size, BASF HiQ-7S19 cm3) to a total weight of 1.0 g. The solid mixture was then loaded into the stainless tube reactor. The reactor was heated in a controlled electrical furnace. Gas mixtures fed into the reactor were obtained by mixing methane (99% purity, Praxair) and air (extra dry, Praxair), metered by mass flow controllers. The reaction mixture leaving the reactor was first completely stripped of water by passing over a desiccant before the methane and carbon dioxide concentrations were analyzed by a gas chromatograph (Shimadzu GC-8A). 2.2. Fluidized-Bed Combustion Tests. A schematic of the reactor system, used previously by Constantineau17 and Johnsen et al.18 for other high-temperature fluidized-bed processes, is shown in Figure 1. The major components include a preheater, a 0.66 m high  0.1 m inner diameter stainless-steel fluidized-bed reactor with an expanded freeboard, a filtration unit, and a gas cooler. A removable stainless-steel gas distributor plate, with 34 drilled 1.2 mm diameter holes on a hexagonal grid, was installed between the preheater and reactor. Methane and air were premixed before being introduced into the lower part of the preheater. This gas passed through the distributor into a mixture of the commercial catalyst and limestone. Three zones of the reactor were heated by electrical furnaces, each of which could be controlled individually. Differential pressure transducers measured the pressure drops between 90 and 280 mm above the distributor. Temperatures and pressures were recorded by a data acquisition system. Gas in or above the bed was sampled through a 6.3 mm stainless-steel tube, with a small plug of quartz wool at its open end to prevent particles from entering. The samples were dried and collected into Teflon bags. The composition of these samples was then determined by the gas chromatograph. The bed materials employed were limestone and commercial catalyst (G-74D), with the same size range as in the kinetic tests. The total mass of bed material was 2.7 kg, with a 40:1 mass ratio of limestone/catalyst. The static bed height was ∼200 mm. At the operating temperature between 723 and 923 K, the limestone acted as inert particles. The minimum fluidization velocity (Figure 2) decreased from 0.045 m/s at

Figure 1. Schematic of the experimental combustion unit.

Figure 2. Minimum fluidization velocity derived from the pressure drop versus superficial velocity at two temperatures. room temperature to 0.038 m/s at 923 K. The superficial velocity was 0.10-0.25 m/s during the experiments.

3. REACTOR MODEL A one-dimensional bubbling bed model was developed to simulate the reactor based on two regions, bubbles and emulsion, with an interphase mass-transfer coefficient, kbe, to characterize the transfer of gas between these two phases. Mole balances are written for the CH4 concentration, Cb,CH4, in the bubble phase and, Ce,CH4, in the emulsion, as functions of height, z  dCb, CH4

-δub

dz

¼ δγb Fkr Cmb, CH4 þ δkbe ðCb, CH4 - Ce, CH4 Þ

-ð1 - δÞumf

976

dCe, CH4 ¼ ð1 - δÞð1 - εmf ÞFkr Cme, CH4 dz - δkbe ðCb, CH4 - Ce, CH4 Þ

ð2Þ

ð3Þ

dx.doi.org/10.1021/ef101573y |Energy Fuels 2011, 25, 975–980

Energy & Fuels

ARTICLE

where kr is the reaction rate constant, m is reaction order, and γb is the void fraction of solids dispersed in the bubbles. From previous experimental data, we expect19,20 γb ≈ 0.01-0.001. Here, we assume an intermediate value, γb = 0.005. The overall gas interchange coefficient between the bubble and emulsion phases, kbe, was obtained from Sit and Grace21   2:0umf Dεmf ub 1=2 kbe ¼ þ 6:77 ð4Þ db db 3 with εmf = 0.52 as the void fraction in the bed at the minimum fluidization. Bubble growth was estimated on the basis of the correlation by Darton et al.22 pffiffiffiffiffi 0:54ðu0 - umf Þ0:4 ðz þ 4 A0 Þ db ¼ ð5Þ g 0:2 -4

Figure 3. Kinetic rate constant for methane combustion kinetics over the G-74D catalyst for a fitted reaction order of 0.61 at 1 vol % CH4 (balance air).

where A0 is the area of the distributor per orifice, 2.3  10 m . The bubble volume fraction in the bubbling regime was estimated from u0 - umf ð6Þ δ¼ ub 2

where ub is the bubble rising velocity through the bed, estimated as ub ¼ u0 - umf þ ubr

ð7Þ

with ubr, the bubble rise velocity relative to the emulsion phase, given by pffiffiffiffiffiffi ubr ¼ 0:711 gdb ð8Þ In these expressions, the bubble phase gas flux, including throughflow, is 

ub ¼ ub þ 3umf

ð9Þ

Figure 4. Effects of the bed temperature on methane conversion for u0 = 0.15 m/s.

Equations 2 and 3 are solved with the boundary conditions Cb, CH4 ¼ Ce, CH4 ¼ CCH4 , in

at z ¼ 0

ð10Þ

assumed to describe the kinetic data:

Gas samples from the bed are assumed to be weighted according to the proportion to the volume fractions of two phases,23 so that Csample ¼ δCb, CH4 þ ð1 - δÞCe, CH4

-RCH4 ¼ k0 expð - E=RTÞCmCH4 ¼ kr CmCH4

ð11Þ

By least-squares fitting the experimental data, the pre-exponential factor, k0, was evaluated to be 8.6  106 mol s-1 kgcatalyst-1 (mol/m3)-0.61, the activation energy, E, was evaluated to be 100.3 kJ/mol, and the reaction order, m, was evaluated to be 0.61. The kinetic rate constant, kr, is reported in Figure 3 as an Arrhenius plot. 4.2. Methane Conversion and Concentration Profile in Fluidized Bed. Figure 4 shows the methane conversion as a function of the bed temperature for methane inlet concentrations of 0.5, 1, and 3 vol %, all at a superficial gas velocity of 0.15 m/s. The methane conversion increased with the temperature for a given inlet methane concentration. From eq 14, the kinetic rate constant increases quickly with the temperature, causing more methane to be consumed with an increasing temperature. Under oxygen-rich conditions, the products should be predominantly carbon dioxide and water, with little carbon monoxide. Lee and Trimm6 and Mouaddib et al.24 confirmed that methane was oxidized to carbon dioxide over Pd supported on alumina with a temperature range of 500-850 K under oxygenrich conditions. As illustrated in Figure 5, the CO concentration increased with the bed temperature. From a fundamental point of

In the freeboard for the conditions studied, there were relatively few particles, so that the methane concentration drawn by the sampling probe can be regarded as the volumetric flow average of the two phases at the upper surface of the dense bed. 

Csample ¼

ð1 - δÞumf δu Ce, CH4 , out þ b Cb, CH4 , out u0 u0

ð14Þ

ð12Þ

The methane conversion at the exit can be expressed by   ð1 - δÞumf δub 0 Ce, CH4 , out þ Cb, CH4 , out ð13Þ XCH4 ¼ 1 u0 u0

4. RESULTS AND DISCUSSION 4.1. Kinetic Evaluation of the Catalyst. With oxygen present in great excess, the following simple kinetic model was 977

dx.doi.org/10.1021/ef101573y |Energy Fuels 2011, 25, 975–980

Energy & Fuels

ARTICLE

Figure 5. CO concentration at the reactor exit for u0 = 0.15 m/s.

Figure 7. Effects of superficial velocity on methane conversion for a bed temperature of 923 K.

Figure 6. Effects of the inlet methane concentration on methane conversions.

view, CO does not seem to be an intermediate of methane oxidation but can be produced by steam reforming of methane modulated by the well-known water-gas shift (WGS) reaction CH4 þ H2 OCO þ 3H2 CO þ H2 O T CO2 þ H2

0 ΔH298 ¼ 206 kJ=mol

ð15Þ

0 ΔH298 ¼ - 41 kJ=mol

ð16Þ

Figure 8. Effects of inlet methane concentration on the methane concentration profile at u0 = 0.15 m/s and a bed temperature of 773 K.

Figure 6 shows the variation in methane conversion as a function of the inlet methane volumetric concentration for a superficial velocity of 0.15 m/s and bed temperatures of 723, 823, and 923 K. The methane conversion is seen to have decreased with an increasing inlet methane concentration. For a bed temperature of 823 K, the conversion decreased from 84.2 to 56.7% as the inlet methane concentration increased from 0.15 to 3 vol %. For a bed temperature of 923 K, the methane conversion always exceeded 85%. The effects of superficial velocity on methane conversion are presented in Figure 7 for a bed temperature of 923 K. The methane conversion decreased with an increasing superficial gas velocity at a given methane concentration. For example, for a methane inlet concentration of 0.5 vol %, the conversion dropped from 98.5 to 86.5% when the superficial velocity increased from 0.1 to 0.25 m/s. This can be explained by larger bubbles at a higher gas superficial velocity, causing the contact time to be shorter and more methane to bypass the bubble phase. Methane axial profiles are plotted in Figure 8. The dimensionless methane concentration decreased gradually with height below the bed surface and increased abruptly above the bed surface. When the sample probe was immersed in the bubbling

From previous studies,6,24 oxidation catalysts are also active catalysts for the methane steam reforming and WGS reactions. The reforming process is strongly endothermic and, therefore, favored by high temperature, while the WGS reaction is weakly exothermic and, therefore, favored by low temperatures. For an oxygen-rich atmosphere, the following reactions also occur: 1 H 2 þ O2 f H2 O 2 1 CO þ O2 f CO2 2

0 ΔH298 ¼ - 241:8 kJ=mol

ð17Þ

0 ΔH298 ¼ - 283 kJ=mol

ð18Þ

Hydrogen was not detected at the reactor exit. It can be seen in Figure 5 that the carbon monoxide concentration was very low compared to the inlet methane concentration, e.g., less than 750 ppm, even at a bed temperature of 923 K and an inlet methane concentration of 3 vol %. Therefore, it is reasonable to ignore the CO in the simulation model. 978

dx.doi.org/10.1021/ef101573y |Energy Fuels 2011, 25, 975–980

Energy & Fuels

ARTICLE

fluidized bed, the composition of the gas drawn by the probe depends upon the fraction of the time spent by the probe in each phase.25 In the freeboard, however, the gases from the individual phases have mixed, so that a sample probe draws gas with a volumetric flow average. Therefore, the transition to the volumetric flow average above the bed surface results in a jump in the methane concentration. 4.3. Comparison of Experimental Data to Model Predictions. Calculations from the two-phase model are compared to the experimental data in Figures 4, 6, 7, and 8. In each case, it can be seen that the model predictions give correct trends and compare reasonably well to the experimental data. Given the modest superficial gas velocities, the reaction in the freeboard was not considered in the simulations. However, small amounts of catalyst splashed into the freeboard, probably accounting for the finding that experimental conversions tended to be slightly higher than predicted values.

5. CONCLUSIONS Methane and air were mixed to investigate combustion of lowconcentration CBM. Combustion tests were carried out in a bubbling fluidized-bed reactor with a mixture of limestone and G-74D catalyst as the bed material. The kinetics of the principal reaction over the catalyst was determined in a fixed-bed microreactor. The derived kinetic parameters were then applied in a two-phase fluidized-bed reactor model. Model predictions were generally in good agreement with the experimental data. Other conclusions of this study were as follows: (1) Both methane conversion and CO concentration increased with an increasing bed temperature. The increasing CO concentration with the bed temperature can be attributed to the methane steam reforming and WGS reactions. (2) The methane conversion decreased when the inlet methane concentration increased from 0.15 to 3 vol %. (3) Less methane was converted when the superficial velocity was augmented from 0.1 to 0.25 m/s, because of the shorter contact time and more methane bypassing in the bubble phase. (4) The sampled methane concentration decreased gradually with height below the bed surface and increased abruptly at the bed surface. This occurred because of the transition from time-average volumetric sampling in the bed to volumetric flow-average sampling in the freeboard.

Cb,CH4,out = methane concentration in the bubble phase out of the bed (mol/m3) Ce,CH4,out = methane concentration in the emulsion phase out of the bed (mol/m3) D = diffusion coefficient (m2/s) E = activation energy (J/mol) db = bubble diameter (m) kr = reaction rate constant [mol s-1 kgcatalyst-1 (mol/m3)-0.61] kbe = mass-transfer coefficient between the bubble and emulsion phase (s-1) k0 = pre-exponential factor [mol s-1 kgcatalyst-1 (mol/m3)-0.61] m = reaction order RCH4 = methane reaction rate (mol s-1 kgcatalyst-1) T = temperature (K) u0 = superfical velocity (m/s) ubr = bubble rise velocity with respect to the emulsion phase (m/s) ub = bubble rising velocity through the bed (m/s) ub* = rise velocity of bubble gas (m/s) umf = minimum fluidized velocity (m/s) XCH4 = experimental methane conversion (%) XCH40 = calculated methane conversion (%) z = axial bed height (m) γb = void fraction of solids dispersed in the bubbles δ = bubble volume fraction in the bed εmf = void fraction in the bed at minimum fluidization F = catalyst density (kg/m3)

’ REFERENCES (1) Su, S.; Beath, A.; Guo, H.; Mallett, C. Prog. Energy Combust. Sci. 2005, 31 (2), 123–170. (2) Yang, Z.; Zhang, L.; Tang, Q. Natural Gas Industry 2010, 30 (2), 115–118 (in Chinese). (3) Iamarino, M.; Chirone, R.; Pirone, R.; Russo, G.; Salatino, P. Combust. Sci. Technol. 2002, 174 (11), 361–375. (4) Trimm, D. L.; Lam, C. W. Chem. Eng. Sci. 1980, 35 (6), 1405– 1413. (5) Choudhary, T. V.; Banerjee, S.; Choudhary, V. R. Appl. Catal., A 2002, 234 (1-2), 1–23. (6) Lee, J. H.; Trimm, D. L. Fuel Process. Technol. 1995, 42 (2-3), 339–359. (7) Chaouki, J.; Guy, C.; Sapundzhiev, C.; Kusohorsky, D. Ind. Eng. Chem. Res. 1994, 33 (12), 2957–2963. (8) Gosiewski, K.; Warmuzinski, K. Chem. Eng. Sci. 2007, 62 (10), 2679–2689. (9) Gosiewski, K. Chem. Eng. J. 2005, 107 (1-3), 19–25. (10) Iamarino, M.; Chirone, R.; Lisi, L.; Pirone, R.; Salatino, P.; Russo, G. Catal. Today 2002, 75 (1-4), 317–324. (11) Iamarino, M.; Salatino, P.; Chirone, R.; Russo, G. Proc. Combust. Inst. 2002, 29 (1), 827–834. (12) Grace, J. R. Chem. Eng. Sci. 1990, 45 (8), 1953–1966. (13) Zukowski, W. Fuel 2000, 79 (14), 1757–1765. (14) Hayhurst, A. N.; John, J. J.; Wazacz, R. J. Symp. (Int.) Combust., [Proc.] 1998, 27 (2), 3111–3118. (15) Foka, M.; Chaouki, J.; Guy, C.; Klvana, D. Chem. Eng. Sci. 1994, 49 (24), 4269–4276. (16) Sotudeh-Gharebaagh, R.; Chaouki, J. Energy Fuels 2007, 21 (4), 2230–2237. (17) Constantineau, J. P. Fluidized bed roasting of zinc sulfide concentrate: Factors affecting the particle size distribution. Ph.D. Dissertation, University of British Columbia, Vancouver, British Columbia, Canada, 2004. (18) Johnsen, K.; Ryu, H. J.; Grace, J. R.; Lim, C. J. Chem. Eng. Sci. 2006, 61 (4), 1195–1202. (19) Fang, F.; Li, Z.; Cai, N. Energy Fuels 2009, 23 (1), 207–216.

’ AUTHOR INFORMATION Corresponding Author

*Fax: þ86-23-65111832. E-mail: [email protected].

’ ACKNOWLEDGMENT We are grateful to the China Scholarship Council for making Mr. Zhongqing Yang’s studies in Canada possible. The Chongqing Science and Technology Commission (CSTC) is also acknowledged for its financial support (Project 2009BA6067). ’ NOMENCLATURE A0 = distributor area per orifice (m2) Csample = concentration in the sampling probe (mol/m3) Cb,CH4 = methane concentration in the bubble phase (mol/m3) Ce,CH4 = methane concentration in the emulsion phase (mol/ m3 ) CCH4,in = inlet methane concentration (mol/m3) 979

dx.doi.org/10.1021/ef101573y |Energy Fuels 2011, 25, 975–980

Energy & Fuels

ARTICLE

(20) Kunii, D.; Levenspiel, O. Fluidization Engineering, 2nd ed.; Butterworth-Heinemann: Boston, MA, 1991. (21) Sit, S. P.; Grace, J. R. Chem. Eng. Sci. 1981, 36 (2), 327–355. (22) Darton, R. C.; LaNauze, R. D.; Davidson, J. F.; Harrison, D. Trans. Inst. Chem. Eng. 1977, 55 (4), 274–280. (23) Davidson, J. F. Inst. Energy 1984, 2 (1), 1–7. (24) Mouaddib, N.; Feumi-Jantou, C.; Garbowski, E.; Primet, M. Appl. Catal., A 1992, 87 (1), 129–144. (25) Grace, J.; Bi, H.; Zhang, Y. Chem. Eng. Sci. 2009, 64 (10), 2522– 2524.

980

dx.doi.org/10.1021/ef101573y |Energy Fuels 2011, 25, 975–980