Investigation of low concentration methane combustion in a fluidized bed with Pd/Al₂O₃ as catalytic particles

Abstract

The catalytic combustion characteristics of low concentration (0.15–3 vol%) methane combustion in a lab-scale fluidized bed with 0.5 wt% Pd/Al₂O₃ as catalytic particles were studied experimentally. A mathematical model was developed according to the flow and reaction characteristics of the fluidized bed. The effects of bed temperature and inlet methane concentration on combustion were investigated and the kinetic characteristics were analyzed. The results show that methane conversion increases with increasing bed temperature, while it decreases slightly as the inlet methane concentration increases. The reaction order was evaluated as 0.608, and the activation energy was determined to be 96 185 J mol⁻¹ during low-concentration methane catalytic combustion in the fluidized bed. It was also found that the reaction in the fluidized bed was controlled by kinetics when the temperature was below 450 °C. When the temperature exceeded 450 °C, the reaction kinetic constant increased with increasing temperature, and the reaction was eventually controlled by kinetics, mass transfer and diffusion. A comparison of the results showed that the values calculated by the mathematical model agreed well with the experimental data.

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1. Introduction

Low concentration methane exists extensively in coal bed methane in coal mines and exhaust gases from chemical production processes.^1,2 Its volumetric concentration changes under different operating conditions, which makes it difficult to utilize. In some cases, low concentration methane emits to the atmosphere directly without treatment. It is a waste of non-renewable energy and severely affects global warming, as the greenhouse effect of methane is 21–23 times greater than that of carbon dioxide.³ Therefore, it is important to find appropriate approaches to utilize low concentration methane resources.

Compared to conventional flame combustion, catalytic combustion has several advantages in burning low concentration methane, such as low light-off temperature, low energy consumption, wide scope of application, high efficiency, low secondary pollution, and high oxidizing speed. Moreover, full oxidation of methane can be achieved with suitable catalysts at low temperatures.^4,5 Furthermore, it has been determined that air–fuel mixtures can be burned effectively by catalytic combustion without regard to flammability limits.^6,7 Thus, catalytic combustion can be an effective way to utilize low concentration methane.

Though the kinetic mechanism of methane catalytic combustion involves multi-step surface reactions, the global combustion reaction can be approximately expressed as follows:

CH₄ + 2O₂ → CO₂ + 2H₂O ΔH⁰₂₉₈ = −802 kJ mol⁻¹ (1)

There are different types of combustion catalysts, but noble metals such as platinum, palladium, and ruthenium are generally accepted as the most active catalysts for low-temperature combustion.^8,24 Trimm and Lam studied the characteristics of methane combustion on platinum-alumina fiber catalysts.⁹ Lee and Trimm summarized the change rules for activation energies and reaction orders by varying the experimental working conditions.¹⁰ The variation of activation energy depends on catalyst and temperature, and the reaction order for methane was found to be in the range of 0.5–1.

Chaouki et al. studied the characteristics of methane catalytic combustion in a cyclic catalytic reactor with reversed flow.¹¹ Pd/Al₂O₃ (0.5%; mass fraction) was employed as catalytic particles in the fluidized bed. It was found that the use of a catalytic bed material contributes to lowering of the NO_x emissions and that there would be no NO_x production from methane catalytic combustion at a temperature of 650 °C, if an appropriate choice of operating parameters was used. Gosiewski et al. studied the combustion characteristics of mine ventilation air in a catalytic flow reversal reactor.^12,13 They showed that the reactor can work only under the condition that the methane concentration remains steady in the mine ventilation air. At higher temperatures, homogeneous combustion occurred in the combustor, causing the formation of local hotspots and deactivation of the catalyst, which made the device very unstable. Several factors could account for the deactivation of the catalyst, but the most critical reason was overheating. Once heat was generated without its effective removal from the catalyst, overheating and thermal deactivation would take place successively.¹⁴

Iamarino et al. studied the characteristics of fluidized-bed combustion using a premixed lean methane–air mixture over a copper-based catalyst.¹⁵ It was found that heat extraction from the catalytic combustion system could be accomplished conveniently in a fluidized-bed reactor. Moreover the fluidized-bed reactor ran isothermally under different operating conditions without the occurrence of local hotspots.¹⁶ Another advantage of fluidized bed combustion is its adaptability to a variety of fuels; thus, it has been used widely for low-grade solid fuel combustion applications, such as inferior coal, coal gangue, sludge, and biomass.^17–19 The authors also investigated a premixed fluidized-bed reactor under lean conditions, with 0.15–3% inlet methane concentrations and using Cu/γ-Al₂O₃ as catalyst particles. It was found that methane could be completely converted when the bed temperature was below 700 °C.¹⁴ Zukowski et al. investigated catalytic combustion of gaseous fuel in a bubbling fluidized bed with manganese oxides as the bed material.²⁰ Heterogeneous reactions on the surfaces of the catalyst particles and homogeneous reactions in the inter-particle spaces were both observed when the bed temperature exceeded 750 °C. Besides heterogeneous methane oxidation taken place on the catalytic surface, and homogenous reactions could be neglected. Hayhurst et al. studied catalytic combustion of propane–air mixtures in a fluidized bed of hot sand, with Pt/Al₂O₃ as catalyst pellets.²¹ Interestingly, sand inhibited the combustion, whereas platinized alumina particles catalyzed the combustion during the experiments. Foka and Sotudeh-Gharebagh investigated the combustion characteristics of premixed natural gas in a turbulent fluidized bed.^22,23 During the experiments, the temperature varied over the range 400–600 °C and the inlet mixture contained 4 vol% methane. It was found that the turbulent flow regime was more suitable for the combustion of natural gas than the bubbling regime. Xu, et al. studied the methane catalytic combustion over noble metal catalysts and Mn_1−xCe_xO_2±y catalysts.^24,25 They found that the reaction order with respect to oxygen decreased with temperature while the order with respect to methane remained constant when lean-burn methane combustion were investigated in a low temperature range of 473–673 K, and 3.3 wt% Pd/Al₂O₃ acted as catalytic particles. And, Pd oxide affected the adsorption and activation of reactants.

However, current research is focused mainly on natural gas, and the catalytic combustion of low concentration methane with concentrations below 3 vol% has been reported rarely. The effects of bed temperature and inlet methane concentration on the combustion characteristics, mathematical models, and kinetic analysis of low concentration methane catalytic combustion in fluidized beds should be further studied. For this study, low concentrations of methane were produced by premixing methane and air, and the catalytic combustion characteristics were studied experimentally in a lab-scale fluidized bed reactor with 0.5 wt% Pd/Al₂O₃ as catalytic particles. A mathematical model of the low concentration methane catalytic reactions in the fluidized bed was established according to the flow and reaction features, and the kinetic characteristics of the reactions were analyzed.

2. Experimental section

The experimental arrangement is shown schematically in Fig. 1. A bubbling fluidized bed reactor was used and it was composed mainly of a pre-heater (1), a gas distributor plate (2), a furnace (3), a sampling tube (4), a cyclone separator (5), and a cooling water pipe (6). The removable gas distributor plate was made of stainless steel and installed between the pre-heater and furnace. It had 34 drilled holes with diameters of 1.2 mm, which were regularly distributed on a hexagonal grid. The furnace consisted of two main parts: a stainless-steel fluidized-bed reactor with a height of 0.66 m and an inner diameter of 0.1 m, and an expanded freeboard with a height of 0.15 m and an inner diameter of 0.37 m. The purpose of the expanded freeboard was to recycle the bed material particles when the fluidized bed ran under the bubbling condition. The height between the reactor and the expanded freeboard was 0.14 m. The two zones of the furnace were heated by electrical heaters that could be controlled independently. During the experiment, three thermocouples, shown as T₁, T₂, and T₃ in Fig. 1, were used to measure temperatures; and three pressure measuring points, shown as of P₁, P₂, and P₃, were installed. T₁ and P₁ were basically located in a horizontal line, which was located 90 mm above the distributor plate; T₂ and P₂ were located 280 mm above the distributor plate. The data measured by the thermocouples and pressure sensors were collected by a data sampling system with a sampling frequency of 10 Hz. Gas was sampled from the outlet of the furnace through a stainless-steel tube 6.3 mm in length. To keep particles from entering into the tube, a small plug of quartz wool was used. After drying, the samples were collected into sampling bags made of Teflon. The composition of the samples was determined by using a gas chromatograph (Shimadzu GC-8A).

Fig. 1 Schematic drawing of the experimental system.

The bed materials employed were inert particles (limestone) and a commercial catalyst (G-74D), with a mass ratio of 40 : 1. During the experiments, the reaction temperature varied within the range of 450–700 °C, which is below the decomposition temperature of limestone; consequently, limestone can be considered as inert particles for low concentration methane combustion.³ Particles of catalyst and limestone were in the same size range of 180–325 μm and were evenly mixed before being loaded into the fluidized bed. The total weight of the bed material was 2.8 kg and the height of the static bed was about 0.2 m. The active component of the commercial catalyst G-74D was 0.5 wt% palladium on an alumina carrier, and the mass of the catalyst was 70 g. During the experiments, the low concentration methane fed into the reactor was obtained by premixing methane (99% purity, Praxair) and air (extra dry, Praxair) so the varying fluidized velocity or gas mixture ratio could be controlled by regulating the flow of air and methane. To ensure the accuracy of the experimental data, the experiments were repeated three times under the same conditions.

The concentration of CO produced was less than 10 ppm during the experiments, which was quite low compared to other products. The methane conversion fraction (X_CH4) was defined as follows:

where C_CH4(in) is the inlet methane concentration (mol m⁻³), and C_CH4(out) is the outlet methane concentration (mol m⁻³).

3. Mathematical model

A mathematical model of the relevant reaction and kinetics was established on the basis of the characteristics of the fluidized bed. The bed materials employed were limestone and commercial catalyst (Geldart B) and the superficial velocity varied from 0.1–0.25 m s⁻¹ in a bubbling fluidized bed during the experiments. It was assumed that bubble and emulsion phases exist in the fluidized bed, the gas in the emulsion phase runs through the bed at the minimum fluidization velocity and the excess gas runs through the bed in the form of bubbles. Thus, the variations of methane in the bubble and emulsion phases can be expressed by:

[the variation of methane in the bubble phase] = [the amount of methane reacted in the bubble phase] + [the amount of methane transferred from the bubble phase to the emulsion phase]

[the variation of methane in the emulsion phase] = [the amount of methane reacted in the emulsion phase] + [the amount of methane transferred from the emulsion phase to the bubble phase]

Therefore, the mathematical expressions for the bubble and emulsion phases are given by:

Studies have shown that heterogeneous reactions dominate the catalytic combustion of methane, as the activity of the catalyst could inhibit the occurrence of homogeneous reactions. Thus, it was reasonable to ignore homogeneous reactions in the mathematical model.

Therefore, the reaction rate of methane in the bubble phase can be expressed by:

R_b = γ_bk_rC_b^m (5)

Moreover, the reaction rate of methane in the emulsion phase can be expressed by:

R_e = (1 − ε_mf)k_rC_e^m (6)

It is worth pointing out that catalytic reactions on the surfaces of particles are continuous and stable for low concentration methane in the emulsion phase. That is to say, the mass of reacted species transferred from the gas phase to particle surfaces is equal to that consumed by catalytic reactions in unit time. So it can be described by the differential equation:

The efficiency factor of catalysts is a function of the Thiele number, and is given by:

where, φ is the Thiele number for a spherical catalyst, which can be expressed by:where, D_e is regarded as the effective diffusion coefficient of the gas.

Introduce dimensionless factors:

The index of mass transfer between the bubble and emulsion phases was estimated as:

The Stanton number of the mass transfer was given by:

The Damkohler number was expressed by:

Combining the above expressions, eqn (3), (4) and (7) can be simplified as follows:

(Θ_e − Θ_s) = ηDaΘ_s^m (16)

The above eqn (14)–(16) can be solved given the following boundary conditions:

when ξ = 0, Θ_b = Θ_e = 1 (17)

4. Results and discussion

The effects of bed temperature and inlet methane concentration on low concentration methane combustion in a fluidized bed reactor were studied experimentally. Numerical calculations were carried out with the mathematical model and the results were compared and discussed.

4.1 Effects of bed temperature on methane conversion

Fig. 2 depicts the effects of bed temperature in the range of 450–700 °C on low concentration methane combustion in a fluidized bed. It can be seen that methane conversion rises quickly with increasing bed temperature, particularly when it is below 550 °C. When the bed temperature was 450 °C, the methane conversion fraction was 61.4% at an inlet methane concentration of 2 vol% and a fluidization velocity of 0.1 m s⁻¹. When it increased to 550 °C, the methane conversion fraction increased to 96.3%, and full conversion was achieved at a bed temperature of 650 °C.

Fig. 2 Effects of temperature on methane conversion at a fluidization velocity of 0.1 m s⁻¹ (a) α_CH4 = 3%, u₀ = 0.1 m s⁻¹ (b) α_CH4 = 2%, u₀ = 0.1 m s⁻¹ (c) α_CH4 = 0.15%, u₀ = 0.1 m s⁻¹ (d) α_CH4 = 0.3%, u₀ = 0.1 m s⁻¹.

The bed materials employed during the experiments were the catalyst and inert particles. It was reasonable to ignore the homogeneous combustion reactions and to consider that only heterogeneous reactions took place during the process of methane reacting with oxygen on the surfaces of solid catalyst during low concentration methane combustion in the fluidized bed.²⁰ The processes of chemical adsorption, chemical reactions, and chemical desorption occurring on the surfaces of particles were regarded as chemical kinetics processes for heterogeneous reactions, while the others were regarded as diffusion process. The reactions of low concentration methane catalytic combustion were controlled by not only catalytic kinetics and diffusion, but also affected by mass transfer between the bubble and emulsion phases. According to the kinetics of methane catalytic reactions, the kinetic constant and reaction rate increased with rising temperature, and simultaneously diffusion was enhanced.

4.2 Effects of inlet methane concentration on methane conversion

Low concentration methane is produced as a by-product of the coal mining process, and methane emissions from coal mines vary in concentration based on different mining operations. Therefore, the effect of inlet methane concentration on conversion was studied with its volumetric fraction varying from 0.15 to 3%, and the results were shown in Fig. 3. It is clear that the methane conversion fraction decreased at higher inlet methane concentrations. At a superficial velocity of 0.1 m s⁻¹ and a bed temperature of 500 °C, the methane conversion fraction decreased from 97% to 88% as the inlet methane concentration increased from 0.15 to 3 vol% for Fig. 3(a). While it displayed similar trends that the methane conversion fraction decreased with increasing inlet methane concentration in Fig. 3(b)–(d). And, the principal analysis of Fig. 3(a)–(d) are as follows.

Fig. 3 Effects of inlet methane concentration on methane conversion under different conditions (a) u₀ = 0.1 m s⁻¹, T = 500 °C (b) u₀ = 0.1 m s⁻¹, T = 550 °C (c) u₀ = 0.15 m s⁻¹, T = 550 °C (d) u₀ = 0.15 m s⁻¹, T = 600 °C.

The CO concentration was very low (less than 10 ppm) compared with the other products. Therefore, it was reasonable to ignore the component of CO, and the methane conversion fraction can be approximately expressed as the ratio between the molar volume of methane reacted and the molar volume of inlet methane (eqn (18)).

where, R_(CH4) represents the methane reaction rate, τ is the contact reaction time; k_r is the reaction kinetic constant, and m is the reaction order.

Methane conversion can be regarded as an approximate function of methane concentration, which can be seen from the above equation. Based on the selection of catalysts for the methane catalytic combustion reaction, the methane combustion reaction order is in the range between 0.5 and 1. For the catalyst (0.5 wt% Pd/Al₂O₃) employed in these experiments, the reaction order m was 0.608. The details of the reaction order fitting will be shown in the next section. It can be seen from eqn (18) that methane conversion decreases with increasing inlet methane concentration, as m is less than 1.

4.3 Kinetic analysis of low Concentration methane catalytic combustion

The reaction order can be obtained by measuring the methane conversion as it varies with inlet concentration at a given bed temperature. Fig. 4 shows the effects of inlet methane concentration on conversion at a bed temperature of 450 °C. It can be seen that conversion is a power function of inlet methane concentration. It was assumed that the reaction order of oxygen was zero for the one-step reaction between methane and oxygen under the condition that an excess of oxygen was always present during the experiments; moreover, the reaction order of methane was evaluated to be 0.608 by least-squares fitting of the experimental data.

Fig. 4 Effects of inlet methane concentration on conversion at 450 °C.

Similarly, the pre-exponential factor and activation energy can be evaluated by measuring methane conversion as it varies with bed temperature at a given inlet concentration, then the reaction kinetic constant can also be obtained. By least-squares fitting of the experimental data, the pre-exponential factor, k₀, was evaluated to be 1.35 × 10¹⁰ mol (s m_catalyst³)⁻¹ (mol m⁻³)^−0.608 and the activation energy, E, was evaluated to be 96 185 J mol⁻¹.

During low concentration methane catalytic combustion, the reaction is controlled by intrinsic kinetics when the temperature is relatively low. As the temperature increases, the reaction gradually becomes controlled by kinetics, mass transfer, and diffusion together. Fig. 5 shows the dimensionless parameters Da and ηDa varying with temperature; η represents the catalyst efficiency factor, of which the value is less than or equal to 1. The Damkohler number Da is the ratio of the chemical reaction rate to the diffusion rate. It can be seen from Fig. 5 that the Da curve is generally in good agreement with the ηDa curve and η is approximately equal to 1 when temperature is below 450 °C. It should be noted that when the temperature exceeds 450 °C, the value of η decreases with increasing temperature, and the deviation value between the Da and ηDa curves gradually increases. The reaction in the fluidized bed is controlled by kinetics when the temperature is no more than 450 °C; while at higher temperatures, the reaction kinetic constant increases, and the reaction is gradually controlled by kinetics, mass transfer and diffusion.

Fig. 5 Variation of dimensionless parameters Da and ηDa with temperature.

Fig. 6 shows the Stanton number St and the mass transfer index of the Mt variation with temperature in the bubble and emulsion phases. The Stanton number represents the mass transfer between catalyst particle surfaces and external gas. It can be seen that the value of St is an order magnitude larger than that of Mt, that is to say, the effects of mass transfer between the outer surfaces of the catalyst particles and the external gas were obviously significant, and were substantially greater than that of the mass transfer between the bubble and emulsion phases during low concentration methane combustion in a fluidized bed. Consequently, it is reasonable to surmise that the gas concentration on the outer surfaces of the catalyst particles was approximately equal to that in the emulsion phase of the fluidized bed.

Fig. 6 Variation of dimensionless parameters St and Mt with temperature.

4.4 Comparisons of model predictions and experimental data

Numerical calculations were carried out by the mathematical model of low concentration methane combustion in a fluidized bed under various operating conditions, which can be seen in Fig. 2 and 3. Many more comparisons of experimental and calculated data are shown in Fig. 7. It can be seen the model calculated values agreed well with the experimental data and the maximum deviation was less than 15%.

Fig. 7 Comparisons of experimental and calculated data.

The major cause of the errors was the strong gas–solid turbulent flow in the fluidized bed. The turbulence enhanced the effects of mass transfer between the gas and catalyst particles. Simultaneously, a small amount of catalyst particles was splashed into the freeboard by the effect of turbulence. In the mathematical model, considering that most of the methane was consumed in the dense zone, the small number of heterogeneous reactions in the freeboard was ignored. Besides, it was assumed for the simplified model that gas ran through the emulsion phase at the minimum fluidization velocity and the rest of the gas ran through the bed in the form of bubbles. However, the bubbles that formed in the bubbling fluidized bed were not strictly spherical and the growth of bubbles occurred with the merging of bubbles, and there was a certain difference between the shapes and growth of bubbles in the model and in the actual fluidized bed.

5. Conclusions

Low concentration methane catalytic combustion characteristics were experimentally studied in a lab-scale fluidized bed reactor that used 0.5 wt% Pd/Al₂O₃ as catalytic particles. A mathematical model of low concentration methane catalytic reactions in a fluidized bed was developed, and the kinetic characteristics were analyzed. Model predictions were generally in good agreement with the experimental data. The major conclusions can be summarized as follows:

(1) Methane conversion increases with increasing bed temperature, while it decreases slightly with increasing inlet methane concentration. The kinetic rate constant increases rapidly with rising temperature, resulting in more methane being consumed. As the bed temperature increased to 650 °C, methane conversion reached approximately 100%.

(2) The mathematical model of low concentration methane catalytic combustion was developed according to the characteristics of flow and catalytic reactions in a fluidized bed. By least-squares fitting of the experimental data, several kinetic parameters were obtained, the reaction order was evaluated to be 0.608 and the activation energy was evaluated to be 96 185 J mol⁻¹ for the reaction.

(3) According to analysis of the dimensionless parameters Da, St, and Mt, the reactions in the fluidized bed were controlled by kinetics when the temperature was 450 °C. When the temperature exceeded 450 °C, as the reaction kinetic constant of methane increases with increasing temperature, the reaction gradually became controlled by kinetics, mass transfer and diffusion together.

Nomenclature

Cb	Methane concentration in the bubble phase (mol m^−3)
Ce	Methane concentration in the emulsion phase (mol m^−3)
CCH4	Methane concentration (mol m^−3)
CCH4(in)	Inlet methane concentration (mol m^−3)
CCH4(out)	Outlet methane concentration (mol m^−3)
Cs	Material concentration on the outer surfaces of catalyst particles (mol m^−3)
De	Effective diffusion coefficient of a gas (m^2 s^−1)
ωb	Dimensionless velocity of a bubble rising through the bed
ωe	Dimensionless minimum fluidized velocity
E	Activation energy (J mol^−1)
RA	Macro-kinetic reaction rate [mol (s^−1 m^−3)]
Ss	External surface area of catalyst particles (m^2)
Vs	Volume of catalyst particles (m^3)
db	Bubble diameter (m)
kc	Intrinsic kinetic reaction rate constant [mol (s mcatalyst^3)^−1 (mol m^−3)^−0.608]
kg	Gas phase mass transfer coefficient (m s^−1)
kr	Reaction rate constant [mol s^−1 kgcatalyst^−1 (mol m^−3)^−0.608]
kbe	Mass-transfer coefficient between the bubble and emulsion phase (s^−1)
k0	Pre-exponential factor [mol s^−1 kgcatalyst^−1 (mol m^−3)^−0.608]
RCH4	Methane reaction rate (mol s^−1 kgcatalyst^−1)
Rb	Methane reaction rate in the bubble phase (mol s^−1 kgcatalyst^−1)
Re	Methane reaction rate in the emulsion phase (mol s^−1 kgcatalyst^−1)
Φ	Dimensionless methane concentration
Θb	Dimensionless methane concentration in the bubble phase
Θe	Dimensionless methane concentration in the emulsion phase
ξ	Dimensionless axial bed height
φ	Thiele number for a spherical catalyst
Mt	The index of mass transfer between the bubble and emulsion phases
Stm	Stanton number of mass transfer
Da	Damkohler number
m	Reaction order
η	The efficiency factor of catalysts
n	The mass of transmitter substance (mol)
T	Temperature (K)
u0	Superficial velocity (m s^−1)
Tbed	Bed temperature (K)
ub	Bubble rising velocity through the bed (m s^−1)
umf	Minimum fluidization velocity (m s^−1)
XCH4	Calculated methane conversion fraction (%)
z	Axial bed height (m)
τ	Contact reaction time
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

Acknowledgements

The authors would like to say thanks for the financial support of the Natural Science Foundation of China with Project no. 51206200, the Fundamental Research Funds for the Central Universities with Project no. CDJZR12140031, and the Visiting Scholar Foundation of Key Laboratory of Low-grade Energy Utilization Technologies and System, MOE of China in Chongqing University (LLEUTS-201301).

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