Fluidized-Bed Methanation: Interaction between Kinetics and Mass

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Fluidized-Bed Methanation: Interaction between Kinetics and Mass Transfer Jan Kopyscinski, Tilman J. Schildhauer,* and Serge M. A. Biollaz General Energy Research Department, Paul Scherrer Institute, 5232 Villigen PSI, Switzerland ABSTRACT: In the present work, the influence of reaction and mass transfer on the fluidized-bed methanation have been investigated by both experiments and modeling. By applying spatially resolved gas concentration and temperature measurements in a bench-scale fluidizedbed reactor, it was shown that most of the reaction proceeds in the first 20 mm while the temperature increases by 74 K in the first 2 mm of the bed. A CO conversion of practically 100% is achieved. The experimental data indicate that the measured gas composition represents mainly the dense phase and that mass transfer between bubble and dense phase is the dominant effect in the upper part of the bed. A fluidized-bed model is proposed based on the two-phase model approach, hydrodynamic correlations from the literature and kinetic parameters previously determined. Although it was not possible to reproduce all measured phenomena within the methanation reactor, this first attempt to model the fluidized bed provides a better understanding of the behavior of the reactor and the reactions.

1. INTRODUCTION The conversion of coal and dry biomass to a clean fuel such as methane, the so-called substitute or synthetic natural gas (SNG), draws much interest because of the longer availability of these resources. SNG is produced from coal and dry biomass via thermochemical processes: gasification followed by gas cleaning, gas conditioning, methanation of the producer gas, and subsequent gas upgrading. This process allows for an easy and cost-effective carbon dioxide (CO2) removal as the separation of a highly concentrated CO2 stream is inherent for the production of SNG in pipeline quality. SNG is a versatile energy carrier that is interchangeable with natural gas (>95% methane, high heating value). The advantages of producing SNG are the high conversion efficiency (65%), the already existing gas distribution infrastructure such as pipelines and the well-established and efficient end-use technologies, e.g., compressed natural gas (CNG) cars, heating, combined heat and power (CHP), power stations. In the methanation of carbon oxides to methane, three independent reactions are important (reactions I-III). 3H2 þ CO T CH4 þ H2 O

Methanation :

ΔHR0 ¼ - 206:28 kJ=mol Water-gas shift :

CO þ H2 O T CO4 þ H2

ΔHR0 ¼ - 41:16 kJ=mol Boudouard :

ðIÞ

ðIIÞ

2CO T C þ CO2

ΔHR0 ¼ - 172:54 kJ=mol

ðIIIÞ

If the stoichiometric ratio of the reactants H2/CO is at least three or more, carbon monoxide (CO) reacts with hydrogen (H2) to produce methane (CH4) and water (H2O), according to reaction I. However, producer gases from biomass and coal gasifiers usually have a H2/CO ratio of 0.3-2, which is too low for a good CO conversion and long catalyst lifetime. By means of the water gas shift r 2010 American Chemical Society

reaction (WGS, reaction II), the H2/CO ratio can be adjusted by converting CO with H2O to CO2 and additional H2. The Boudouard reaction (reaction III) is also important in systems where the H2/CO ratio is low. On one hand, carbon on the catalyst surface can be considered as a necessary intermediate during the methanation reaction. On the other hand, it may lead to catalyst deactivation if C atoms are not hydrogenated fast enough and form polymeric or graphitic carbon deposits.1 The reaction mechanisms and kinetics have been studied intensively,2-15 since Sabatier and Senderens16 found in 1902, that nickel and other metals catalyze this reaction. In the past 50 years, a few dozen papers have been published regarding the kinetics of CO and CO2 methanation over different nickel catalysts. In all of these studies, only the gas composition at the reactor outlet was measured for different experimental conditions. The rate of the methanation was directly calculated from the conversion of CO or the exit gas concentration of CH4. Recently, by applying spatially resolved gas concentration measurements in a catalytic plate reactor, the effects of reactants (H2, CO) and products (CH4, H2O, CO2) on the rates were analyzed for similar conditions and for the same catalyst as that used in this work. Furthermore, the kinetic parameters of the proposed Langmuir-Hinshelwood rate expressions of the methanation and WGS reaction were determined.17 There is no consensus in the literature on the elemental steps for the methanation of CO on a nickel surface. Two mechanisms have been proposed with different rate expressions. The common assumption is that the methanation proceeds via an intermediate carbon species formed either by direct CO dissociation (leading to a adsorbed C species)7-10,18-20 or by dissociation of an oxygenated carbon compound (COHx complex).4,5,13,21 The methanation reaction is exothermic and, because of the high inlet CO partial pressures in the methanation reactors for Special Issue: IMCCRE 2010 Received: March 15, 2010 Accepted: May 3, 2010 Revised: April 30, 2010 Published: June 9, 2010 2781

dx.doi.org/10.1021/ie100629k | Ind. Eng. Chem. Res. 2011, 50, 2781–2790

Industrial & Engineering Chemistry Research

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Model CP 4900; detector TCD, Columns MSA5, PPU). The complete setup is LabVIEW controlled. Besides the measurement of the axial gas concentration profiles, the fluidized-bed setup allows one to take catalyst samples. 2.2. Experimental Procedure. For the experiments, 100 g of a commercial Ni/γ-Al2O3 catalyst were used (50 wt % Ni/Al2O3, BET surface area = 183 m2/g). The reactor temperature was kept at 320 °C ((5 K). The pressure was 1.3 barabs. The inlet gas mixture consisted of 60 vol % H2, 20 vol % CO, and 20 vol % N2 with a total volume flow of 10 LN/min. Before the experiments, the catalyst was reduced inside the fluidized-bed reactor. During reduction, the volume flow of N2 was constant, while the H2 concentration was increased stepwise from 5 vol % to 70 vol %. The experiment consisted of six runs, where the sampling probe was moved upward and downward. The μ-GC was calibrated with calibration mixtures ((0.5%-2% relative uncertainty) at three to five concentration levels per gas species (H2/CO/CO2/CH4/N2). At the start of each run, the sampling probe was immersed with a back flush of N2 to prevent particles from blocking the probe, and moved down to the porous plate (level zero). The N2 back flush then was stopped and the pump and the μ-GC were started. In the following, the sampling probe was moved up and down automatically by the linear motor (precision of (0.125 mm).

Figure 1. Simplified scheme of the bench-scale fluidized-bed setup.

SNG production, an enormous amount of heat must be removed. In this case, gas-solid fluidized-bed reactors have been successfully applied for the production of SNG from biomass and coalderived synthesis gas in several projects.22-25 Recently, by in situ measurements of gas concentration and temperature profiles in such a fluidized-bed reactor, it could be shown that hydrodynamic and chemical boundary conditions influence the process and, thus, the performance of the reactor.26,27 The objective of this work is to understand more in detail the different processes inside the fluidized-bed methanation reactor based on both experiments and modeling. The measured and calculated gas concentration profiles along the reactor are compared and discussed.

2. EXPERIMENTAL SECTION 2.1. Experimental Setup. The bench-scale fluidized-bed setup is presented in Figure 1. The reactant gases H2, CO, and N2 are fed via calibrated mass flow controllers (Bronkhorst EL-Flow) into a 52-mm-inner-diameter fluidized-bed reactor and distributed by a nonreactive porous metal plate. The temperature of the reactor is measured by a sidewaysimmersed fixed thermocouple ∼5 mm above the porous metal plate (T-bed) and controlled by cooling the outside reactor wall with air. The product gas leaving the reactor is sent through a filter to remove the catalyst fines. Water produced by the methanation reaction is condensed and weighed. The core of the setup is a moveable sampling tube with an outer diameter of 2 mm and a porous titanium plate at the tip. The tube is equipped with a thermocouple (T-probe) at the tip to measure the bed temperature profile. Driven by a linear motor, this probe is immersed axially from the top of the reactor into the catalyst bed. Spacers at two different heights ensure the radial position of the probe. By means of a gas pump, a continuous flow of 5-10 mL/min is transported to the micro gas chromatograph (μ-GC) via an MgSO4 adsorber for removal of the moisture. Gaseous products are analyzed online by a μ-GC with two columns (Varian

3. MATHEMATICAL MODEL Besides the experiments, a first attempt was made to describe the experimental data using a fluidized-bed model. The modeling of such a reactor is advantageous, because it allows one to easily study the influence of the operating conditions and therefore gain a better understanding of the measured phenomena. The presented model is based on the two-phase fluidized-bed model,28,29 to which several improvements were added to account for different assumptions as described below. • Steady-state conditions and ideal gas behavior are assumed. • The bubble phase is solid-free; thus, the reactions are occurring only in the dense phase. • The gas concentration in the dense phase is the same as that inside the catalyst particles. Hence, no laminar boundary layer around the particles is assumed, and possible influence of pore diffusion is neglected. • Plug flow in the bubble and dense phases is applied; thus, no axial dispersion is considered. • Radial gas concentration differences are neglected. • The pressure loss due to the weight of the solid is neglected, because it is small, compared to the absolute pressure. • The reaction system considers only the methanation and the WGS reaction. Kinetic studies have shown that the hydrogenation of an adsorbed carbon intermediate is the rate-determining step.7,10,17,30 The formation of this intermediate carbon species via CO or COH dissociation is very fast. Furthermore, it is assumed that all carbon on the surface is converted to CH4 or CO2. Because of this fact, carbon on the catalyst is not modeled. • The gas flow rate through the dense phase is constant, but not at minimum fluidization condition as assumed by the original two-phase theory.28 Measurements by Werther31,32 showed that less gas is transported in the form of bubbles than is assumed by the original two-phase theory. That means that more gas flows through the emulsion phase. • Because of volume contraction inherent to the methanation in the dense phase, an extra convective mass transfer from the bubble phase into the dense phase must be considered; this is called bulk flow. 2782

dx.doi.org/10.1021/ie100629k |Ind. Eng. Chem. Res. 2011, 50, 2781–2790


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Figure 2. Scheme of the two-phase fluidized-bed methanation model.

With this assumption, the gas flow rate in the dense phase is kept constant. This bulk flow is assumed to proceed immediately at the height at which the molar reduction occurs. • The temperature profile from the bench-scale fluidized bed experiments is used as an input parameter. No energy balance is to be solved. Furthermore, it is assumed that the temperature difference between the bubble phase and the dense phase is negligible. Figure 2 shows the scheme of the one-dimensional homogeneous two-phase model. The total molar gas flow enters the reactor in the bottom and splits into bubble phase flow and dense phase flow. Both phases are connected via mass transfer, which is described by the mass-transfer coefficient and the concentration difference. Furthermore, the convective mass transfer term called bulk flow takes into account the volume contraction in the dense phase, as discussed above. 3.1. Mass Balances. The reactor model considers six gaseous compounds for the methanation and WGS reaction (i = 1, ..., 6): H2, CO, CO2, CH4, H2O, and N2. Experiments and gas concentration measurements proved that higher hydrocarbons such as C2H4, C2H6, and C3H8 are not formed. Based on the scheme in Figure 2 and the model assumptions, the molar balances for the bubble phase can be written as d_nb, i - KG, i aAðcb, i - ce, i Þ - N_ vc xb, i ð1Þ 0 ¼ dh and for the dense phase as

quantity 1 - εb the volume fraction of the dense phase, and the quantity 1 - εmf the volume fraction of the particles. The total bulk flow from the bubble to the dense phase N_ vc is described as the sum of the molar losses due to the reaction and mass transfer in the dense phase. X n_ vc ¼ N_ vc ¼ KG, i aAðcb, i - ce, i Þ dh i X þ ð1 - εb Þð1 - εmf ÞFP A Ri ð3Þ i

The boundary conditions at h = 0 (inlet) are n_ b, i jh ¼ 0 ¼ n_ b, i, feed n_ e, i jh ¼ 0 ¼ n_ e, i, feed

ð4Þ ð5Þ

The molar fractions in the bubble phase and dense phase are equal under the inlet conditions. The molar flow of the bubble phase and the dense phase are calculated via the visible bubble flow defined by Hilligardt and Werther:32,33 Avb ¼ Aψðu - umf Þ ð6Þ where ψ describes the deviation from the original two-phase theory (ψ = 0.69). The visible bubble flow velocity vb, the superficial gas velocity u, and the minimum fluidization velocity umf are based on the empty tube diameter. The terms Au and Avb represent the total volume flow through the reactor and the bubble phase volume flow, respectively. The overall reaction term is defined as X νij rj ð7Þ Ri ¼ To calculate the rates for the species H2, CO, CH4, CO2, H2O, and N2, the following relations are applied: ð8Þ RH2 ¼ - 3r1 þ r2

Here, a is the specific mass-transfer area, A the cross-sectional area of the reactor, KG,i the mass-transfer coefficient, xb,i the molar fraction in the bubble phase, FP the particle density, the

RCO ¼ - r1 - r2

ð9Þ

RCO2 ¼ r2

ð10Þ

RCH4 ¼ r1

ð11Þ

R H2 O ¼ r 1 - r 2

ð12Þ

R N2 ¼ 0

ð13Þ

Here, r 1 and r2 are the reaction rates (given in units of mol kgcat-1 s-1) of the methanation and WGS reaction, respectively. 2783



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Table 1. Hydrodynamic Correlations Used for the Fluidized-Bed Model parameter

unit

correlation and equation 1=3

bubble diameter

m

db ¼ db0 ½1 þ 27:2ðu - umf Þ

theoretical initial bubble diameter

m

db0 = 0.00853

visible bubble flow

m/s

vb ¼ ψðu - umf Þ

bubble velocity

m/s

ub ¼ ψðu - umf Þ þ 0:711υ

m-1

εb ¼ uvbb a ¼ 6εdbb

bubble phase holdup specific mass-transfer area mass-transfer coefficient

m/s

minimum fluidization velocity

m/s

KG, i ¼ umf ¼

umf 3

32, 33

pffiffiffiffiffiffi gdb with ψ = 0.69 and υ = 0.63

32, 33

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4Di, mix εmf ub þ

dP ðFP - Fg Þ εmf 3 φP g 1 - εmf 150η

38

ð15Þ

Ki ¼ Ki, Tref

34, 36

37

where an Arrhenius-type dependency was assumed, according to EA , j Tref kj ¼ kj, Tref exp 1ð16Þ RTref T ΔHi Tref exp 1RTref T

ð1 þ 6:84hÞ

πdb

3.2. Hydrodynamics. The fluidized-bed model mainly follows the two-phase theory. The numerical model describes the hydrodynamic parameters by applying correlations for the height-dependent bubble size (db), the bubble velocity (ub), the bubble gas holdup (εb), the specific surface area (a), the mass-transfer coefficient (KG,i), and the minimum fluidization velocity (umf), as summarized in Table 1. The correlations proposed by Werther and Hilligardt31-36 are used to calculate the visible bubble flow, the bubble diameter, and the bubble velocity along the fluidized-bed height. The mass transfer between the bubble phase and the dense phase is described by applying the widely used correlation by Sit and Grace.37 The minimum fluidization velocity is calculated via the simplified Ergun equation,38 where dP is the particle size, η the viscosity of the gas mixture, φP the sphericity of the catalyst particles, FP the particle density, Fg the gas density, and εmf the bed voidage under minimum fluidization conditions. 3.3. Reaction Kinetics. The reaction term in the dense phase (eq 2) includes the rate equations and kinetic parameters as determined by Kopyscinski et al.17 The proposed kinetic model for the methanation reaction assumes the hydrogenation of an adsorbed carbon, C* þ H* f CH* þ *, as the rate-determining step (RDS). The RDS for the WGS reaction is assumed to be CO* þ OH* T CO2* þ H*. The rate equation of the methane production is described by the following Langmuir-Hinshelwood expression: k1 KC pCO 0:5 pH2 0:5 r1 ¼ ð14Þ 2 1 þ KC pCO 0:5 þ KOH pH2 O pH2 - 0:5

The rate equation of the WGS reaction is described as k2 KR pCO pH2 O - pCO2 pH2 =Keq r2 ¼ 2 pH2 0:5 1 þ KC pCO 0:5 þ KOH pH2 O pH2 - 0:5

ref(s) 1:21

ð17Þ

2

2

Table 2. Kinetic Data Used for the Fluidized-Bed Model of the Methanation Reactora parameter

a

value -1

k1,Tref EA,1

1.16 mol s kgcat-1 bar-0.5 74.1 kJ/mol

k2,Tref

7.42 mol s-1 kgcat-1 bar-1.5

EA,2

161.6 kJ/mol

KC,Tref

1.77 bar-0.5

ΔHC

-61.0 kJ/mol

KOH,Tref

0.66 bar-0.5

ΔHOH

-72.3 kJ/mol

KR,Tref ΔHR

0.34 -6.7 kJ/mol

Data taken from ref 17.

they were calculated at each bed height. The heat capacity was used to calculate the equilibrium constant Keq of the WGS reaction, which is described by the following van’t Hoff equation: ∂lnðKP Þ ΔHR0 ¼ ð18Þ ∂T RT 2 The standard reaction enthalpy can be computed for each temperature by Z T νi cp, i dT ð19Þ ΔHR ðTÞ ¼ ΔHR0 ðT0 Þ þ T0

The reference temperature T0 under standard conditions is 298.15 K (25.0 °C). The correlation for the heat capacity is shown in eq 20, and the corresponding coefficients A, B, C, D, and E were taken from the DIPPR Project 801 database.39 2 2 C=T E=T cp, i ¼ A þ B þD ð20Þ sinhðC=TÞ coshðE=TÞ The molecular diffusion coefficient of a species in a gas mixture was calculated according to eq 21 (from Wilke et al.40). The corresponding binary diffusion coefficients were calculated according to eq 22 (from Fuller et al.41,42). 1 - xi Di, mix ¼ P ðxj =Di, j Þ

with Tref = 598.15 K. The values of the kinetic parameters of both reactions are given in Table 2. 3.4. Thermodynamic Gas Properties. The thermodynamic gas properties (diffusion coefficient, gas viscosity, and heat capacity) vary with temperature and gas composition; therefore,

ð21Þ

j

(

T 1:75 ½ð1=Mi Þ þ ð1=Mi Þ1=2 Di, j ¼ 0:01013 P P p½ð vi Þ1=3 þ ð vj Þ1=3 2 2784

) ð22Þ



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Table 3. Results of the Fluidized-Bed Experiments: Conversion of CO and H2; Yield of CH4 and CO2; Deviation of the Atomic Balance of C, H, and O; Condensed H2O; and Theoretical and Real Outlet Flow theoretical outlet

Conversion

Yield

Deviation of the Atomic Balance

H2O

real outlet

run

flow (LN/min)

XCO (%)

XH2 (%)

YCH4 (%)

YCO2 (%)

ΔC (%)

ΔHb

ΔOb

(g/min)c

flow (LN/min)d

1

6.2

100

89.8

89.9

6.8

3.3

1.4

1.1

1.22

6.09

2

6.2

100

89.5

88.9

7.0

4.1

1.8

0.7

1.22

6.09

3 4

6.2 6.2

100 100

89.1 89.9

88.3 85.9

7.0 6.6

4.7 7.5

1.8 4.3

0.7 1.6

1.22 1.22

6.10 6.00

5

6.2

100

89.2

89.4

7.3

3.3

1.2

0.0

1.22

6.12

6

6.2

100

89.3

89.8

7.2

2.9

0.9

0.2

1.22

6.13

L

6.2

100

89.5

88.7

7.0

4.3

1.9

0.7

1.22

6.1

a

a

Calculated with software HSC Chemistry 5.1 (equilibrium). bAtomic balance was calculated using the rate of the condensed water. c Condensed reaction water; weight increase was measured with a balance. d Real outlet flow of the experiments.

Here, x is the molar fraction, T the temperature (given in Kelvin), p the pressure (givenP in units of Pa), M the molecular weight (given in units g/mol), and v the atomic diffusion volume (given in units of m3/mol). The dynamic viscosity of the gas mixture was computed according Wilke43 (see eq 23). X xi η P i ð23Þ ηm ¼ ðxj φji Þ i j

with

0:5 0:25 2 1 þ ηi =ηj Mj =Mi φi, j ¼ 0:5 8 1 þ Mi =Mj

The gas viscosities of the pure components were determined via the DIPPR39 correlation (see eq 24). AT B ηi ¼ ð24Þ 1 þ ðC=T Þ þ ðD=T 2 Þ All the equations described above lead to a nonlinear initialvalue problem containing ordinary differential equations (mass balances) and a set of algebraic equations (e.g., hydrodynamics, kinetics, thermodynamic properties). The DDAPLUS algorithm within the Athena package developed by Stewart and Caracotsios44,45 solves the differential-algebraic equation system.

4. RESULTS AND DISCUSSION 4.1. Experimental Results. Mass Balances and Conversion. To characterize the performance, the conversion of CO and H2, the yield of CH4 and CO2, and the deviation of the atomic balances of C, H, and O are calculated for each run and averaged for the experiment. Moreover, the theoretical and real outlet volume flow is determined. The results are summarized in Table 3. It is worth mentioning that the cooled and condensed product gas stream had a temperature of ∼20 °C ((1 K). Thus, the reaction water was not completely condensed. The corresponding saturation pressure of water (2300 ( 200 Pa) leads to a water concentration in the product gas of ∼2.3 vol %. The atomic hydrogen and oxygen balance are calculated using the measured rate of water condensation and further considering the saturation pressure of water in the product gas stream.

The conversion of CO approached 100% for all runs, and the conversion of H2 was ∼89.5%. The CH4 yield was 88.7%, and the CO2 yield was 7.0%. The atomic balances for C, H, and O were almost even. For carbon, the deviation in the atomic balance was mostly in the range of 2.9%-4.7%; only once was it as high as 7.1%. The hydrogen balance deviated from 100% by 0.9%-1.9%; only once was it as high as 4.3%. The atomic balance for oxygen deviated by a maximum of 1.6%. The theoretical outlet flow was determined by means of equilibrium calculation, using the software tool HSC Chemistry 5.1. The real outlet flow was determined using the measured N2 concentration in the freeboard as an internal standard, considering the condensed water and the saturation pressure of water in the condensed product gas. The actual outlet flows correspond well to the theoretically calculated outlet flows. A volume reduction of up to 38% is calculated and measured. Gas Species Concentration and Temperature Profiles. Figure 3 shows the measured axial profiles of the dry gas species composition and the temperature. In this diagram, the dry gas concentration of N2, H2, CH4, CO, and CO2 (in vol %) and the bed temperature (in °C) are plotted on the x-axis, whereas the reactor height (given in millimeters) is plotted on the y-axis. In this experiment, the H2 concentration decreases rapidly from 60 vol % before the gas distributor to 9.9 vol % after 12.5 mm. The H2 concentration then increases slightly until the end of the bed to 12.3 vol %. The CH4 concentration rises at the same time to a maximum value of ∼45.4 vol % after 22.5 mm and declines then to 41.2 vol % at the end of the bed. The CO concentration drops nearly instantaneously in the first 2 mm to practically zero. The N2 concentration grows rapidly and reaches its maximum of 45.3 vol % after 9 mm, because of very fast reaction and volume contraction. Thereafter, the N2 concentration decreases and passes through a local minimum of 40.6 vol % after 22.5 mm and grows again slightly until the end of the bed. CO2 is formed in the first 2 mm (8 vol %), passes then through a local minimum at 9 mm (2.2 vol %), and increases again to value of 3.4 vol %. Until the end of the bed, the CO2 concentration falls slightly to 3.2 vol %. Directly on the gas distributor (height = 0), a CH4 concentration of 25 vol % and a H2 concentration of 30 vol % are measured. In addition, a temperature increase of 74 K, compared to the set bed temperature of 320 °C, is measured. The set temperature of 320 °C is reached after 10 mm and remains constant until the end of the bed. 2785


Industrial & Engineering Chemistry Research The profile shows a clear step in the concentrations of H2, CH4, and N2 at a height of ∼95 mm. At this height, the end of the catalyst bed is expected (illustrated by a wave-line), which is confirmed by temperature measurements. The sudden change of measured concentrations between bed and freeboard indicates that the measured gas concentrations represent mainly the dense phase and that significant concentration differences between the dense phase and the bubble phase exist. The high initial rate and the CO2 production can be explained by the high temperatures of 394 to 380 °C in the first 4 mm. At this temperature range, the methanation rate (CO and H2

Figure 3. Measured axial profiles of dry gas species concentration and bed temperature during methanation reaction with 100 gcat for an inlet flow of 10 LN/min (H2:CO:N2 = 6:2:2) and a bed temperature of Tbed = 320 °C: N2 (denoted by symbol ), H2 (dark diamond symbol), CH4 (gray-shaded diamond symbol), CO (gray-shaded circle symbol), CO2 (dark circle symbol), and Tbed (denoted by symbol þ).

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conversion) is fast, but also the WGS reaction plays an important role, which results in the formation of CO2, as was shown in the kinetic experiments in the catalytic plate reactor.17 As soon as the CO concentration is below the equilibrium concentration of the WGS reaction, CO2 is converted via the reverse water gas shift reaction. Thus, the CO2 concentration decreases. From 4 mm to 15 mm, the rate of CH4 formation and H2 conversion becomes slower. One reason could be limitations in mass transfer. More precisely, in the first 3 mm, all of the CO from the dense phase is consumed; thus, CO must be delivered from the bubble phase into the dense phase, where the catalyst is present. This CO mass transfer might be slower than the intrinsic rate of CO disappearance caused by adsorption and hydrogenation. Hence, the methanation rate is solely dependent on the mass transfer of CO from the bubble phase to the dense phase. The decrease of the CH4 concentration and the increase of the H2 concentration in the upper part of the bed are the result of the mass transfer between the dense phase and the bubble phase, which is explained as follows. One must notice that CH4 is only produced in the dense phase where the catalyst particles are fluidized. Since the bubble phase can be assumed to be particlefree, the increase in CH4 concentration is solely caused by the mass transfer from the dense phase. That means that the concentration of CH4 in the dense phase is always higher than the concentration of CH4 in the bubble phase. In the upper part of the bed, the mass transfer of CH4 from the dense phase into the bubble phase exceeds the production of CH4 in the dense phase and by this, the CH4 concentration in the dense phase is decreasing. Most of the CO and, therefore, most of the adsorbed carbon is converted in the first 20 mm of the catalyst bed. At the end of the bed, both the dense phase and the bubble phase are mixed and the resulting CH4 concentration in the freeboard is somehow between the two. For H2, the explanation is the same, but here, the H2 concentration in the bubble phase is always higher than in the dense phase. In the upper part of the bed, more H2 is transported from the bubble phase as it is consumed in the dense phase. As mentioned earlier, the concentration of N2 increases, then passes through a maximum and decreases. The concentration of

Figure 4. Measured (symbols) and calculated (lines) dry axial gas species composition profiles for 100 gcat at 320 °C for an inlet flow of 10 LN/min with 60% H2, 20% CO, and 20% N2. Bed height is ∼95 mm. 2786



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Figure 5. Calculated hydrodynamic data along the fluidized bed: (a) bubble diameter, (b) specific surface area and bubble gas holdup, and (c) bubble rise and superficial gas velocity.

N2 then passes through a local minimum and increases slightly again until the end of the bed. The increase in the first 9 mm is explained by the volume contraction during the methanation reaction; therefore, it can be assumed that, in this region, the reaction is the dominant effect and that the mass transfer between the bubble phase and the dense phase is almost negligible. Then, the opposite occurs: the N2 concentration passes through a maximum and decreases. This can be explained by mass transfer between the two phases. H2 and CO are consumed in the dense phase; therefore, the concentrations of these species are larger in the bubble phase. The N2 concentration in the bubble phase is smaller than in the dense phase. That means that H2 and CO are transferred from the bubble phase into the dense phase and that N2 is transported from the dense phase into the bubble phase, because of the concentration differences. Above 15 mm, the N2 concentration passes through a local minimum and increases again. This might suggest that N2 is now transferred from the bubble phase into the dense phase. That means the N2 concentration in the

bubble phase in this part of the bed must be larger than in the dense phase. This can be confirmed by the concentration step at the end of the catalyst bed: the N2 concentration in the freeboard is greater than that in the dense phase. In summary, in the bottom region of the bed, carbon (most probably adsorbed carbon atoms, Cads) is deposited because of the dissociation of CO. CO2 is formed via WGS reaction, especially in hotspot regions. The Cads atoms are hydrogenated to methane. In the upper part of the bed, the CH4 concentration decreases, whereas the concentrations of both H2 and N2 increase. This can be explained by the mass transfer between the bubble phase and the dense phase. At the end of the bed, both phases are mixed and the freeboard gas concentrations are somewhere between the concentrations of the bubble phase and the dense phase. It can be concluded that, using this technique, mainly the dense phase of the fluidized bed is probed. It is worth mentioning that without the high spatial resolution within the first 20 mm, the important characteristics of this system cannot be observed. 2787


Industrial & Engineering Chemistry Research The spatially resolved gas concentration and temperature measurement is a very suitable tool to gain detailed information about catalytic reaction systems in fluidized-bed reactors. 4.2. Modeling Results. It was experimentally shown that the measured gas concentration represents mainly the gas composition in the dense phase. Because of this fact, the calculated dry axial gas species concentrations of the dense phase can be directly compared to the measured concentrations, as shown in Figure 4. In this diagram, the dry axial gas composition (expressed in terms of vol %) is plotted versus the height of the fluidized bed (distance above the gas distributor). The symbols in the figure represent the measured gas volume fractions, and the lines in the figure represent the calculated gas compositions. The fluidized bed ends at a height of 95 mm. The fast initial rate in the first few millimeters is represented accurately by the two-phase model. Here, the chemical reaction is dominating, and the mass transfer between the bubble phase and the dense phase is almost negligible. The largest discrepancy between the model results and the experimental data are in the zone between 4 mm to 10 mm. Here, the model shows that the mass transfer is almost equal to the reaction term, which leads to plateaus in the H2 and CH4 concentration profiles. However, the measured gas concentration profiles indicate that the reaction is still the dominating process until a bed height of 10 mm. The deviation may be explained by the uncertain validity of modeling assumptions applied (e.g., extra bulk flow, bubble gas holdup) and of the hydrodynamic correlations used (e.g., bubble size, mass transfer). The hydrodynamic correlations are based on measurements at ambient temperature and pressure in larger fluidized beds without reaction. Thus, the correlations may not be valid for small bench-scale reactors. At the end of the bed, the model predicts correctly the gas concentrations. The three diagrams of Figure 5 show the bubble diameter db (Figure 5a), the specific surface area a and the bubble gas holdup εb (Figure 5b), and the bubble rise velocity ub and the superficial gas velocity u (Figure 5c), all relative to the height of the fluidized bed. The profile of the bubble size versus the height of the fluidized bed shows a slight decrease in the bubble diameter from 14 mm to 13 mm in the entrance region. Then, above a bed height of 10 mm, the bubble size increases almost linearly, up to a value of 21 mm. This is equivalent to 40% of the bed diameter. The slow descent of the bubble size is explained by the decrease of the superficial gas velocity (see Figure 5c). The correlation proposed by Werther34 indicates that the bubble size is proportional to u - umf. The fast decrease of the superficial gas velocity in the first 10 mm of the fluidized bed confirms that the main H2 and CO conversion and volume reduction happen in that region. From the 10 mm-mark to the end of the bed, the superficial gas velocity decreases only slightly. The specific surface area decreases from 150 m2/m3 to 40 m2/m3 along the bed, because of an increasing bubble size and a decreasing bubble gas holdup. The latter decreases from εb = 0.35 at the beginning of the bed to εb = 0.12 at the end of the fluidized bed.

5. CONCLUSIONS A deeper insight into the processes within the fluidized-bed methanation reactor was gained by dedicated experiments and by modeling. By means of a movable sampling probe, the axial gas species concentration and temperature profiles along the height of the fluidized bed were measured. By applying this technique, it

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was found that (1) the dense phase of the fluidized bed is probed; (2) the fluidized bed can be divided into a CO-rich part (entrance region) and CO-lean part (upper part of the bed); (3) the main conversion occurs within the first 20 mm (very fast reaction); and (4) the mass transfer is dominating in the upper part of the bed. A two-phase fluidized-bed model was developed containing hydrodynamic correlations from the literature and kinetic parameters previously determined for the same catalyst. The results of the fluidized-bed model show that (1) the initial slope of the gas composition profiles and the outlet composition are in good agreement with the experimental data points; and (2) within the range of 4-40 mm, the calculated and measured gas concentration profiles differ. The reason might be the hydrodynamic correlations used, which are based on measurements at ambient temperature and pressure in larger fluidized beds without reaction. Thus, the correlations may not be valid for small bench-scale reactors. A sensitivity study and an attempt to modify model assumptions and hydrodynamic correlations from literature are the subject of further investigation and model improvement.

’ AUTHOR INFORMATION Corresponding Author

*E-mail: [email protected].

’ ACKNOWLEDGMENT The financial support of Novatlantis is gratefully acknowledged. The authors thank P. Hottinger, T. Marti, and J. Schneebeli for technical assistance. ’ NOTATION Parameters

A = cross-sectional area of the reactor (m2) a = specific mass transfer area (m2/m3) ci = concentration of species i (mol/m3) cp,i = specific heat (kJ K-1 mol-1) db0 = initial bubble diameter (m) db = bubble diameter (m) dP = particle diameter (m) dt = reactor diameter (m) Di,j = binary diffusion coefficient of species i and j (m2/s) Di,mix = molecular diffusion coefficient of species i (m2/s) EA = activation energy (kJ/mol) g = acceleration due to gravity; g = 9.80665 m/s2 h = height or distance from the gas distribution (m) ΔHR = heat of reaction (kJ/mol) ΔHi = heat of adsorption (kJ/mol) KG,i = mass-transfer coefficient of species i (m/s) Keq = equilibrium constant of the water-gas shift reaction Ki = adsorption constant of species i (units differ; see Table 2) Ki,Tref = pre-exponential factor for adsorption constant Ki (units differ; see Table 2) kj = reaction constant of reaction j (units differ; see Table 2) kj,Tref = pre-exponential factor for rate coefficient kj (units differ; see Table 2) Mi = molar weight of species i (g/mol) m = mass (kg) N_ vc = total bulk flow from the bubble to the dense phase (mol s-1 m-1) n_ i = molar flow of species i (mol/s) 2788


Industrial & Engineering Chemistry Research pi = partial pressure of species i (bar or Pa) R = gas constant; R = 8.314 J mol-1 K-1 rj = rate of reaction for eqs 1 (methanation) and 2 (water gas shift) (mol kgcat-1 s-1) Ri = rate of disappearance or formation of species i (mol kg cat-1 s -1 ) T = temperature (K or °C) u = superficial gas velocity based on tube diameter (m/s) u b = bubble velocity (m/s) u mf = minimum fluidization gas velocity based on tube diameter (m/s) v b = visible bubble phase velocity based on tube diameter (m/s) vi = atomic diffusion volume (m3/mol) xi = molar fraction of species i Xi = conversion of species i Yi = yield of species i Greek symbols

ε = void fraction or porosity η = gas viscosity (Pa s) νij = stoichiometric factor of species i in reaction j F = density (kg/m3) υ = parameter for the bubble rise velocity φP = sphericity of the catalyst particle ψ = parameter for the visible bubble flow Subscripts and Superscripts

b = bubble phase cat = catalyst E = emulsion or dense phase G = gas mf = minimum fluidization condition P = particle ref = reference tot = total

’ REFERENCES (1) Bartholomew, C. H. Carbon deposition in steam reforming and methanation. Catal. Rev.;Sci. Eng. 1982, 24, 67–117. (2) Mills, G. A.; Steffgen, F. W. Catalytic Methanation. Catal. Rev. 1974, 8, 159–210. (3) Ponec, V. Some Aspects of the Mechanism of Methanation and Fischer-Tropsch Synthesis. Catal. Rev.;Sci. Eng. 1978, 18, 151–171. (4) Vannice, M. A. The Catalytic Synthesis of Hydrocarbons from Carbon Monoxide and Hydrogen. Catal. Rev. 1976, 14, 153–191. (5) Van Herwijnen, T.; Van Doesburg, H.; De Jong, W. A. Kinetics of the methanation of CO and CO2 on a nickel catalyst. J. Catal. 1973, 28, 391–402. (6) Ross, J. R. H. Metal catalysed methanation and steam reforming. In Catalysis; Royal Society of Chemistry: London, 1985; Vol. 7, pp 1-45. (7) Klose, J.; Baerns, M. Kinetics of the methanation of carbon monoxide on an alumina-supported nickel catalyst. J. Catal. 1984, 85, 105–116. (8) Ho, S. V.; Harriott, P. The kinetics of methanation on nickel catalysts. J. Catal. 1980, 64, 272–283. (9) van Meerten, R. Z. C.; Vollenbroek, J. G.; de Croon, M. H. J. M.; van Nisselrooy, P. F. M. T.; Coenen, J. W. E. The kinetics and mechanism of the methanation of carbon monoxide on a nickel-silica catalyst. Appl. Catal. 1982, 3, 29–56. (10) Inoue, H.; Funakoshi, M. Kinetics of Methanation of Carbon Monoxide and Carbon Dioxide. J. Chem. Eng. Jpn. 1984, 17, 602–610. (11) Hayes, R. E.; Thomas, W. J.; Hayes, K. E. A study of the nickelcatalyzed methanation reaction. J. Catal. 1985, 92, 312–326.

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(12) Kai, T.; Furusaki, S.; Yamamoto, K. Methanation of Carbon Monoxide by a Fluidized Catalyst Bed. J. Chem. Eng. Jpn. 1984, 17, 280–285. (13) Ibraeva, Z. A.; Nekrasov, N. V.; Yakerson, V. I.; Gudkov, B. S.; Golosman, E. Z.; Beisembaeva, Z. T.; Kiperman, S. L. Kinetics of methanation of carbon monoxide on a nickel catalyst. Kinet. Catal. 1987, 28, 386. (14) Chen, H.; Adesina, A. A. Kinetic modelling of methanation reaction over a Co-Mo/SiO2 catalyst. J. Chem. Technol. Biotechnol. 1994, 60, 103–113. (15) Sehested, J.; Dahl, S.; Jacobsen, J.; Rostrup-Nielsen, J. R. Methanation of CO over Nickel Mechanism and Kinetics at High H2/CO Ratios. J. Phys. Chem. B 2005, 109, 2432–2438. (16) Sabatier, P.; Senderens, J. B. New methane synthesis. Acad. Sci. 1902, 314, 514–516. (17) Kopyscinski, J.; Schildhauer, T. J.; Vogel, F.; Biollaz, S. M.; Wokaun, A. Applying spatially resolved concentration and temperature measurements in a catalytic plate reactor for the kinetic study of COhydrogenation. J. Catal. 2010, 271, 262–279. (18) Araki, M.; Ponec, V. Methanation of carbon monoxide on nickel and nickel-copper alloys. J. Catal. 1976, 44, 439–448. (19) Happel, J.; Cheh, H. Y.; Otarod, M.; Ozawa, S.; Severdia, A. J.; Yoshida, T.; Fthenakis, V. Multiple isotope tracing of methanation over nickel catalyst: II. Deuteromethanes tracing. J. Catal. 1982, 75, 314–328. (20) Alstrup, I. On the kinetics of CO methanation on nickel surfaces. J. Catal. 1995, 151, 216–225. (21) Coenen, J. W. E.; van Nisselrooy, P. F. M. T.; de Croon, M. H. J. M.; van Dooren, P. F. H. A.; van Meerten, R. Z. C. The dynamics of methanation of carbon monoxide on nickel catalysts. Appl. Catal. 1986, 25, 1–8. (22) Kopyscinski, J.; Schildhauer, T. J.; Biollaz, S. M. Production of synthetic natural gas (SNG) from coal and dry biomass;A technology review from 1950 to 2009. Fuel 2010, DOI: 10.1016/j.fuel.2010.01.027. (23) Greyson, M.; Demeter, J. J.; Schlesinger, M. D.; Johnson, G. E.; Jonakin, J.; Myers, J. W. Synthesis of Methane, Technical report, Bureau of Mines, Report of Investigation 5137; U.S. Department of the Interior: Washington, DC, 1955. (24) Streeter, R. C. In Recent Developments in Fluidized-Bed Methanation Research; 9th Synthetic Pipeline Gas Symposium; Bituminous Coal Research, Inc.: Monroeville, PA, 1977; pp 153-165. (25) Hedden, K.; Anderlohr, A.; Becker, J.; Zeeb, H.-P.; Cheng, Y.-H. Gleichzeitige Konvertierung und Methanisierung von CO-reichen Gasen, Technologische Forschung und Entwicklung;Nichtnukleare Energietechnik; DVGW-Forschungsstelle Engler-Bunte-Institut, Universit€at Karlsruhe: Karlsruhe, Germany, 1986. (Prepared for Bundesministerium f€ur Forschung und Technologie, Forschungsbericht T 86-044, Universit€at Karlsruhe.) (26) Seemann, M. C.; Schildhauer, T. J.; Biollaz, S. M. A.; Stucki, S.; Wokaun, A. The regenerative effect of catalyst fluidization under methanation conditions. Appl. Catal., A 2006, 313, 14–21. (27) Kopyscinski, J.; Schildhauer, T. J.; Biollaz, S. M. A. Employing Catalyst Fluidization to Enable Carbon Management in the Synthetic Natural Gas Production from Biomass. Chem. Eng. Technol. 2009, 32, 343–347. (28) May, W. Fluidized-bed reactor studies. Chem. Eng. Prog. 1959, 55, 49–56. (29) van Deemter, J. J. Mixing and contacting in gas-solid fluidized beds. Chem. Eng. Sci. 1961, 13, 143–154. (30) Schoubye, P. Methanation of CO on some Ni catalysts. J. Catal. 1969, 14, 238–246. (31) Werther, J. Der Einfluss des Anstr€ombodens auf die Str€ omungsmechanik von Gas/Feststoff-Wirbelschichten. Chem. Ing. Tech. 1977, 49, 901and MS 533/77. (32) Hilligardt, K.; Werther, J. Local bubble gas hold-up and expansion of gas/solid fluidized beds. Ger. Chem. Eng. 1986, 9, 215–221. (33) Hilligardt, K.; Werther, J. Lokaler Blasengas-Holdup und Expansionsverhalten von Gas/Feststoff-Wirbelschichten. Chem. Ing. Tech. 1985, 57, 622–623and MS 1377/85. 2789



ARTICLE

(34) Werther, J. Die Bedeutung der Blasenkoaleszenz f€ur die Auslegung von Gas/Feststoff-Wirbelschichten. Chem. Ing. Tech. 1976, 48, 339and MS 349/76. (35) Werther, J. Zur Problematik der Massstabsvergr€osserung von Wirbelschichtreaktoren. Chem. Ing. Tech. 1977, 49, 777–785. (36) Werther, J. In Proceedings of the 4th International Conference on Fluidization, Kashikojima, Japan, 1983; Kunii, D., Toei, R., Eds.; United Engineering Trustees: New York, 1984; pp 93-102. (37) Sit, S. P.; Grace, J. R. Effect of bubble interaction on interphase mass transfer in gas fluidized beds. Chem. Eng. Sci. 1981, 36, 327–335. (38) Kunii, D.; Levenspiel, O. Fluidization Engineering, 2nd Ed.; Butterworth-Heinemann: Boston, 1991. (39) Design Institute for Physical Properties. DIPPR Project 801 Full Version , 2005. (40) Wilke, C. R.; Lee, C. Y. Estimation of Diffusion Coefficients for Gases and Vapors. Ind. Eng. Chem. 1955, 47, 1253–1257. (41) Fuller, E.; Schettler, P.; Giddings, J. C. New method for prediction of binary gas-phase diffusion coefficients. Ind. Eng. Chem. 1966, 58, 18–27. (42) Fuller, E. N.; Ensley, K.; Giddings, J. C. Diffusion of halogenated hydrocarbons in helium. The effect of structure on collision cross sections. J. Phys. Chem. 1969, 73, 3679–3685. (43) Wilke, C. R. A Viscosity Equation for Gas Mixtures. J. Chem. Phys. 1950, 18, 517–519. (44) Stewart, W. E.; Caracotsios, M. Computer-Aided Modeling of Reactive Systems, 1st Ed.; Wiley-Interscience: Hoboken, NJ, 2008. (45) Stewart, W. E.; Caracotsios, M. Athena Visual Studio, www. athenavisual.com.

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