Ind. Eng. Chem. Res. 2004, 43, 4721-4731
4721
Compact Radial Reactor with a Structured Porous Metal Catalyst for the Conversion of Natural Gas to Synthesis Gas: Experiment and Modeling V. A. Kirillov,* A. S. Bobrin, N. A. Kuzin, V. A. Kuzmin, A. B. Shigarov, V. B. Skomorokhov, E. I. Smirnov, and V. A. Sobyanin Boreskov Institute of Catalysis, Prospekt Akademika Lavrentieva, 5, 630090 Novosibirsk, Russia
To perform the catalytic conversion of natural gas to synthesis gas, a compact radial reactor supplied with a structured porous metal catalyst has been designed. The reactor was tested at normal pressure using natural gas and air. The reactor can operate without preheating of the inlet gas mixture. A novel structured porous metal catalyst was developed on the basis of the 7% Ni/R-Al2O3 commercial catalyst and Ni-Cr powders. The catalyst is characterized by a low hydraulic pressure drop and a high tortuosity coefficient of the regular structure. The coefficient of heat conductivity of the structured catalyst and the coefficient of mass transfer between a gas flow and the catalyst were determined in special experiments. The catalyst temperature and outlet gas concentrations were experimentally investigated with respect to the gas-air flow rate and methane concentration in the mixture. The catalyst temperature drop along the reactor radius was no higher than 230 °C. The maximum catalyst temperature does not exceed 1090 °C for all runs. At volumetric flow rates of the gas mixture of 4000-40 000 h-1 and O2/CH4 ratios of 0.6-0.9, methane is completely converted, and the maximum space time yield (STY) of the reactor is about 1 L of CH4 per second per liter of reactor. A mathematical model for this process was developed. The results of modeling agree well with the experimental data. One possible application of the developed syngas generator in internal combustion engines is discussed. 1. Introduction The production of synthesis gas (a mixture of CO and H2) from methane or other hydrocarbons is the initial step of a great number of chemical processes. In principle, syngas can be produced by the steam reforming, partial oxidation, or autothermal reforming of methane. Each approach has its advantages and disadvantages. The catalytic steam reforming of methane, which is the technique most widely used in industry for many years, is a highly power- and capital-consuming process. The recent decade has seen a serious effort aimed at developing the catalytic partial oxidation of methane
CH4 + 1/2O2 ) CO + 2H2
∆H°298 ) -36 kJ/mol (1)
The weakly exothermic “dry” reaction 1 permits one to use adiabatic reactors, to avoid water evaporation, and to provide a favorable outlet syngas composition of H2/ CO ) 2. All of these factors reduce the capital costs by a factor of 1.5 as compared to steam reforming.1 Despite the fact that the catalytic partial oxidation of methane (reaction 1) was discovered in 1946,2 it was overlooked for a long time. Probably, this was due to the explosive properties of methane-air mixtures, catalyst coking in the dry atmosphere, and the appearance of a severe hot spot at the entrance of the fixed catalyst bed. Nevertheless, the potential of partial oxidation has led, in the past decade, to the initiation of of efforts toward the development and testing of different catalysts and * To whom correspondence should be addressed. Tel./Fax: (007) 3832 341187. E-mail:
[email protected].
supports that are able to provide reaction 1 to its thermodynamic equilibrium and avoid coking of the catalyst without the addition of steam. The mechanism of the catalytic partial oxidation of methane has not yet been investigated in detail. This fact can be explained by the existence of a complex reaction network. The apparent mechanism of partial methane oxidation was discussed in some previous studies,3-6 and is assumed to consist of CH4 combustion (reaction 2), CH4 steam reforming (reaction 3), and the water-gas shift reaction (reaction 4)
CH4 + 2O2 ) CO2 + 2H2O
∆H°298 ) -803 kJ/mol (2)
CH4 + H2O a CO + 3H2
∆H°298 ) +206 kJ/mol (3)
CO + H2O a CO2 + H2
∆H°298 ) -41 kJ/mol (4)
Moreover, for the Ni/R-Al2O3 catalyst,4 the active component can exist in different states, such as NiAl2O4, NiO, and Ni, along the catalyst bed depending on the O2 concentration in the gas flow. The above-mentioned factors explain the fact that the literature lacks general and reliable kinetic expressions for the partial oxidation of methane. In a few works dealing with the mathematical modeling of reactors designed for this reaction,5-7 the authors had to use kinetic data for reactions 2-4 that were obtained for different catalysts and, as a rule, for different gas flow compositions. On the other hand, the development of an appropriate technology requires a solution of the following nontrivial problems: design of a catalyst, investigation of the reaction kinetics, and design of a reactor providing the
10.1021/ie030785s CCC: $27.50 © 2004 American Chemical Society Published on Web 05/04/2004
4722 Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004
thermal mode required for the partial oxidation of methane. This is associated with the probable formation of “hot spots” resulting in the deactivation or sintering of the catalyst. To prevent such effects, it is necessary to design a thermally stable and heat-conducting catalyst, as well as a novel type of reactor that is characterized by compact construction, high efficiency, and safe operation in the catalytic partial oxidation of methane to synthesis gas. The interesting concepts based on heat coupling between the endothermic and exothermic catalytic reactions (steam reforming and combustion of methane) in a structured reactor are presented in refs 8 and 9. Recently,10 a similar concept was applied to the design of a compact planar reactor producing hydrogen for fuel cells. At the same time, we have developed11 a compact catalytic radial-type reactor supplied with an internal water heat exchanger and a specially designed structured porous metal Pt-containing catalyst for the complete catalytic oxidation of methane. The present paper focuses on the further development of this approach to the conversion of natural gas and the experimental and modeling investigations of the novel reactor operation modes. 2. Novel Structured Catalyst for Natural Gas Conversion Analysis of the literature data shows that the main trend in the development of catalysts for the partial oxidation of methane is associated with supported nickel catalysts, involving the addition of oxides of alkaline earth metals, chromium, zirconium, etc., prepared by different methods, such as coprecipitation of components, mixing of powders or pastes with subsequent molding or pressing, and precipitation of the support with salts of the active component.3,12-14 The commercial nickel catalyst GIAP-315 (7% Ni/R-Al2O3) was chosen on the development of a structured porous metal catalyst for the partial methane oxidation. The preparation technology of this catalyst involves the following operation sequence: preparation of a mixture using GIAP-3 and a nickel powder, preparation of a reinforced steel net, deposition of the mixture on the net and its drying, and manufacturing of flat and corrugated strips for the formation of catalytically active channels. The next step of sintering of the strips by themselves and with a gas-distribution tube of the synthesis gas generator is performed in an inert medium. The prepared catalysts are shown in Figure 1. The as-prepared porous metal catalyst has a specific surface area of SBET ) 3.4 m2/g and its main pore volume is formed by pores ranging from 15 to 100 µm in radius (Table 1). The catalyst composition (in mass %) is 12.5%(7% Ni/Al2O3) + 85.5% Ni + 2% Cr. Scanning electron microscopy (Figure 2a,b) confirmed the large-pore structure of the porous metal catalyst. The porous structure contains two different morphological formations: large rounded conglomerates 8-15 µm in size, formed by the coalescence of particles ∼5 µm in size and smaller aggregates (2-3 µm) formed by particles 0.3-0.5 µm in size. Large pores form cavities between large conglomerates. The catalyst has the structural elements of the initial nickel powder: large round nickel conglomerates that form a metal grid (arrow A in Figure 2a and b). Particles of the nickel-alumina catalyst are distributed in the metal grid (Figure 2a). The prevailing sizes of the supported nickel particles in the GIAP-3 catalyst and the catalyst
Figure 1. General view of the syngas generator, a gas-distribution tube, flat and corrugated strips of the reinforced porous metal catalyst.
layer were 100-200 and 200-300 Å, respectively. One can clearly see nickel particles superimposed on the alumina (arrow C). Therefore, the particles undergo sintering via the surface-diffusion mechanism at a temperature (known as the Tamman temperature) that is considerably lower than the usual sintering temperature of a bulk material. One might propose that heating of the porous metal catalyst samples at 760 °C during the process of preparation provides sintering of the metal nickel powder via the surface-diffusion mechanism, which is accompanied by accretion of the particles at contact points and results in the formation of a monolith. 3. Radial-type Reactor Design for the Conversion of Natural Gas In the development of a reactor for the conversion of natural gas to syngas, the following demands were taken into account: (i) enhanced heat conductivity of the catalyst bed [λs ≈ 1-5 W/(m K)], (ii) highly active and thermally stable catalyst with a long operation life (no less than 6000-8000 h), and (iii) acceptable pressure drop (no more than 1 kPa). Figure 1 presents the general view of a reactor for the conversion of natural gas into synthesis gas that satisfies the above requirements. The reactor is designed as a cylinder that encloses a gas-distribution device, formed from a perforated tube plugged at one end, that is used to feed a methane-air mixture. The structured catalyst bed is formed of flat and corrugated porous catalyst strips wound around the gas-distribution tube and sintered with it. The strips are arranged so that the winds of the following layer overlap the winds of the preceding layer. This results in the formation of a structure consisting of radial and longitudinal channels. The catalyst bed characteristics are a radial thickness of 32 mm (with an inner radius of 22.5 mm and an outer radius of 54.5 mm) and an axial length of 180 mm. The strip characteristics are as follows: the thickness is 0.75 mm, the width is 25 mm, the period of corrugation is 9 mm, the cross section of the longitudinal channel is formed by a trapezoid (with bases of 5 and 2.5 mm and a height of 2 mm), the length of one channel is 10 mm, the width of the radial channel (the gap between two neighboring strips in the layer) is 5 mm, and the reactor volume taking the gas-distribution tube into account is 1600 cm3.
Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004 4723 Table 1. Catalyst Propertiesa VΣ (cm3/g)
catalyst
SBET. (m2/g)
total
GIAP-3 (7 wt % Ni/R-Al2O3) catalyst layer with GIAP-3
6.3 3.4
0.20 0.51
r
105
Å
0.03 0.31
TEM
SEM
SNic (m2/g)
100-200 200-300
250 320
1.50 0.44
a Catalyst characteristics are for the catalyst layer distant from the reinforcing wire mesh. b Size of nickel particles is calculated from pulsed oxygen chemisorption data using the formula dNi ) 104SNi/SFNi (Å), where SNi is the specific surface area of supported nickel, m2/g; S is the specific surface measured by oxygen chemisorption, m2/g; and FNi ) 8.9 g/cm3 is the density of metal nickel. For composite catalysts, SNi is determined from the difference in oxygen adsorption on the catalyst and on the metal support. c Surface area of the metal nickel is calculated taking into account the oxygen adsorption on the metal nickel only.
packing method, the heat conductivity of the gas medium passed through the porous space, and the geometry of the layer. Traditionally, the heat conductivity of a fixed catalyst bed is described as the sum of the heat conductivity of the catalyst bed of the static medium and a convective component caused by gas flow motions16
λs ) λ0 + λconv
(5)
To measure heat conductivity, the standard stationary method was used.16 The meaning of this method is as follows: Assume that one has a hollow cylindrical layer, R1 < r < R2. On the interior surface (r ) R1), there is a heat source with a constant heat flux of q. The temperature of the lateral surface r ) R2 is kept constant and equals T0. For the stationary mode, the dependence of temperature of the hollow cylindrical layer along radius will be described by
T(r) ) T0 +
Figure 2. Photomicrograph of the GIAP-3-based porous metal catalyst (×2000).
During reactor startup, the catalyst bed is preheated with an electric heating unit or a flame burner. As the temperature inside the catalyst bed rises to 600-700 °C, heating is stopped, and the natural gas-air mixture (gas/air excess η ) 0.28-0.47) is directed to the gasdistribution tube. To provide safe operation of the reactor (i.e., to prevent penetration of the flame into the gas-distribution tube during the startup procedure), the diameter of the holes in the gas-distribution tube was 1.5 mm. 4. Heat-Mass Transfer and Hydraulic Resistance Because the structure of the prepared catalyst layer is not traditional for chemical reactors, it was worthwhile investigating the heat conductivity, heat- and mass-transfer processes and hydraulic resistance in the layer to elucidate the prospects of commercialization of the syngas generator. 4.1. Heat Conductivity of the Catalyst Bed. Heat conductivity of the catalyst bed depends on the heat conductivity of the porous metal catalyst itself, the
( )
qR1 R2 ln λ r
(6)
Using a plot of T(r) - T0 versus the inverse radius on a semilogarithmic scale, one can determine the coefficient of heat conductivity of the radial layer λ in static (λ ) λ0) or gas-air mixture flow (λ ) λs) conditions from the slope of the line. During each run, the temperature was monitored in the middle cross section of the layer at three points (with radial coordinates of 34.5, 44.5, and 54.5 mm) at several values of the specific heat flux on the interior layer surface. For the static system, the heat fluxes were q1 ) 7.43 and q2 ) 29.7 kW/m2. As a result, at q1 ) 7.43 kW/m2, the coefficient λ0 is equal to 0.87 W/(m K) and the temperature variation in the catalyst bed is in the range 200-250 °C. At q2 ) 29.7 kW/m2, the heat conductivity is given by λ0 ) 1.37 W/(m K) and the temperature in the catalyst bed is 500-600 °C. The difference in the value of λ0 is associated with the difference in layer temperatures. For the case of air flow through the catalyst bed, the coefficient of heat conductivity was similarly determined but at varied values of the air flow rate and heat flux power. Assuming that the temperature difference between the solid phase and air is minor, we considered the radial layer as a homogeneous medium. The value of λ calculated from eq 6 was taken as the coefficient of effective heat conductivity of the radial layer. The convective component of the heat conductivity coefficient, λconv, can be calculated from eq 5, where λ0 was determined from the experimental data. There is good reason to compare the dependence of λconv on the rate and characteristics of the gas flow with literature data on convective heat conductivity of in a fixed bed.16 For this purpose, in Figure 3, the experi-
4724 Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004
Figure 4. Comparison of the experimental data with reference correlations17,18 for a triangular channel.
Figure 3. Coefficient of radial heat conductivity versus Peλg. Points, experimental data; solid line, correlation in eq 7. Heat flux on the interior layer surface (r ) R1) q ) 29.7 kW/m2.
mental data are presented in terms of the dependence of λs on the product of the Peclet number and the heat conductivity of the gas flow. As shown in Figure 3, the convective part of radial heat conductivity is a linear function of the Peclet number
λconv ) K(Peλgas) ) K(udFcg)
(7)
For the commercial catalysts now in use, the proportionality constant K for a fixed bed, depending on the size and shape of the particles, ranges from 0.15 to 0.3.16 As follows from Figure 3, K ) 0.17 for the structured radial layer, which agrees well with the fixed bed value. However, the total value of the effective coefficient of heat conductivity for the radial layer is much higher than that for a fixed catalyst layer because of the higher heat conductivity of the structured layer under static conditions. 4.2. Mass Transfer in the Catalyst Layer. Coefficients of mass transfer in the porous metal catalyst channels were determined using a model reaction of hydrogen oxidation in the external diffusion region. The method described in section 2 was used to prepare a porous metal catalyst containing 72 wt % Ni, 18 wt % Al, and 10 wt % of platinum catalyst 0.6% Pt/Al2O3. The reaction of hydrogen oxidation with air was performed in a vertical flat gap (8 mm wide and 180 mm long), bounded by side heat-exchanging plates in which cold water was circulating to cool the walls. A catalyst layer was placed between the plates and divided into three equal sections 25 mm in height that were successively arranged. The gaps between the sections were 50 mm in width. Each section of the catalyst layer was formed by a set of corrugated (three pieces) and flat (four pieces) strips of the porous metal catalyst that were sintered by themselves and with the walls of the heat exchanger. The as-prepared catalyst layer contained a set of longitudinal (vertical) channels of equal length (25 mm) that were triangle-shaped (with a base of 2.5 mm and a height of 1.5 mm). Note that the channels did not intersect with one another in the range of a single section. During experimental runs, temperature of the reaction mixture (thermocouples were placed in nine places along the long side of the section base) and the hydrogen concentration (samples from three points were analyzed by a chromatograph and averaged) were measured at the inlet and outlet of each section of the
catalyst layer. Rotameters were used to measure the flow rates of air and hydrogen. All measurements were performed as the system reached its stationary temperature mode. In the processing of the experimental results and the calculation of the Reynolds, Schmidt, and Sherwood numbers, the properties of the airhydrogen mixture (density, viscosity, and diffusion coefficient) were calculated at a temperature determined by averaging the temperature distribution and taking into account the temperatures of the walls bounding the catalyst layer. The runs were performed so that hydrogen was oxidized in the external diffusion mode. Then, the coefficient of mass transfer and the Sherwood number were calculated as follows:
β)-
(
)
[H2,out] Q , ln S [H2,in]
Sh )
βd D
(8)
Figure 4 shows the Sherwood number as a function of the Reynolds number and a comparison with the correlations of calculations for the triangular channels.17,18 In these calculations, we used channels with a hydraulic diameter of 1.2 mm and a length of 25 mm. In Figure 4, the upper line represents calculation results from17
d 0.45 Sh ) 2.7 1 + 0.095Re Sc L
(
)
(9)
The lower line corresponds to the correlation given in ref 18
d 1/3 Sh ) 1.5 Re Sc , L
(
)
d if Re Sc > 7 L
(10)
Figure 4 suggests that our experimental data agree well with the correlations for a triangular channel17,18 in the Reynolds number range from 600 to 1200. 4.3. Hydraulic Resistance. In the experiments designed to study hydraulic resistance, we used the synthesis gas generator described in section 3. Its hydraulic resistance is composed of three main structural units: a gas-distribution tube, a catalytic layer, and an outlet grid. The pressure drop was measured with an alcohol micromanometer at air flow rates varying from 700 to 4000 ncm3/s. Measurements were performed on both the separate units and the whole generator. The hydraulic pressure drop of the catalyst layer makes the main contribution to the overall pressure drop, which is about 75% of the overall pressure drop in the synthesis gas generator. Note that this
Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004 4725 Table 2. Results of Testing of the Synthesis Gas Generator flow rate catalyst of CH4-air air mixture excess temperature (°C) [CH4]out {CH4]out STYa (ncm3/s) η Tin Tout (vol %) [CH4]in (s-1)
Figure 5. Hydraulic resistance of the radial catalyst bed. A comparison of correlation in eq 11 with reference data.19,20
relation between the values of pressure drop was almost constant throughout the whole range of air flow rates. Experimental values of the pressure drop on the catalyst layer are given in terms of the coefficient of hydraulic resistance, σ ) 2∆P/Fu2, as a function of the Reynolds number. For the catalyst layer, the hydraulic diameter of the channel (2.46 mm) and the calculated flow rate of gas in the channels of the outlet corrugated layer were taken as specific parameters in calculations of the Reynolds number. The as-interpreted experimental results are presented in Figure 5. The results are satisfactorily described by the empirical correlation
σ)
6150 + 50 Re
(11)
In Figure 5, the coefficient of resistance of the catalyst layer calculated from the correlations19,20 is compared to the experimental values. As follows from the figure, the experimental data for the radial catalyst layer agree well with the literature correlations when the Reynolds number is higher than 100. 5. Experimental Results for Natural Gas Conversion The aim of the experiments was to investigate the operating mode of the synthesis gas generator and to determine conditions providing complete conversion of natural gas. For this purpose, we experimentally determined the impact of the coefficient of air excess, η, and of the temperature distribution along the catalyst layer on the methane concentration at the catalyst layer outlet. Two thermocouples embedded in the catalyst layer were used to measure the inlet and outlet temperatures of the catalyst at the catalyst layer during the experiments. The flow rates of natural gas and air were measured by rotameters. The validity of these measurements was controlled with a chromatograph that measured the concentrations of methane in the initial mixture after mixing of the natural gas and air. The concentrations of the mixture components at the catalyst layer outlet were also measured with a chromatograph (the mixture was first dried). All measurements were performed as the reactor reached its stationary mode. The flow rate of natural gas was varied from 200 to 1700 ncm3/s. The typical composition of natural gas used in the experiments was C1 ) 96.60 vol %, C2 )1.089 vol %, C3 ) 1.34 vol %, C4 ) 0.44 vol %,
1064 1147 1434 1560 1615 2060 2614 3173 3376 3419 3480 3545 3698 4094 4810 5257 5260 6172 7178 8295 8891 a
0.427 0.446 0.397 0.431 0.442 0.402 0.430 0.435 0.387 0.398 0.410 0.416 0.438 0.433 0.441 0.461 0.431 0.425 0.435 0.420 0.430
1020 1092 996 1051 1091 1073 1090 1090 998 975 1017 1040 1080 1050 1020 1048 990 955 925 855 770
785 815 795 841 880 900 960 1000 908 900 943 973 1020 965 1019 1060 1050 1055 1060 1070 1070
0.09 0.07 0.14 0.05 0.05 0.10 0 0 0.03 0.02 0 0 0 0 0 0 0 0 0 0 0
0.0053 0.0040 0.0078 0.0028 0.0029 0.0057 0 0 0.0017 0.0010 0 0 0 0 0 0 0 0 0 0 0
0.12 0.12 0.18 0.18 0.18 0.24 0.30 0.36 0.42 0.42 0.42 0.42 0.42 0.48 0.54 0.55 0.59 0.71 0.85 0.97 1.03
Space time yield.
Table 3. Composition of Conversion Products at the Synthesis Gas Generator Outlet flow rate of natural gas, (ncm3/s) flow rate of air (ncm3/s) coefficient of air excess, η Tin (°C) Tout (°C)
300 1292 0.44 990 800
958 0.33 852 621
675 815 0.28 810 576
2893 0.44 1010 979
2391 0.36 908 870
1939 0.29 861 715
concentrations of species in outlet dry gas mixture (vol %) H2 25.4 29.0 30.8 26.2 31.0 29.7 CO 12.8 14.3 14.1 12.9 14.7 16.1 CO2 4.24 4.57 3.82 4.42 2.99 2.90 CH4 0.01 1.27 4.83 0.00 1.07 2.13 N2 57.6 50.9 46.5 56.5 50.2 49.2
and i-C4 ) 1.108 vol %. The results are listed in Tables 2 and 3. The last column in Table 2 presents the space time yield (STY), which is the capacity with respect to the natural gas flow rate related to the volume of the radial reactor (volume of catalyst bed + distribution tube ) 1.68 L). The experimental data suggest that the maximum value of STY is about 1 L of methane per secondecond. Along with the above experiments, we analyzed the total composition of the conversion products at the reactor outlet (Table 3). The concentration of methane in the synthesis gas at the reactor outlet depends on the coefficient of air excess, the mixture flow rate, and the temperature in the catalyst layer. Figure 6 shows the impact of the air excess and gas mixture flow rate on the relative concentrations of methane at the reactor outlet. The upper line in Figure 6 corresponds to 10641431 ncm3/s, the middle line to 1560-2060 ncm3/s, and the lower line to 2614-8891 ncm3/s. For our reactor, methane is completely converted if the temperature at the catalyst layer outlet is higher than 940 °C and η is g0.43. It should be noted that the composition of the reaction products at the catalyst layer outlet (see Table 3) is close to the equilibrium composition calculated for the outlet temperature.
4726 Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004
∂Ts )0 ∂r
r ) R1: r ) R2:
Figure 6. Impact of air excess and gas mixture flow rate on the relative methane concentration at the outlet of the synthesis gas generator.
λs
(15)
∂Ts bσrTs4 ) ∂r 1+N
(16)
The boundary condition in eq 16 takes into account the impact of the gas-permeable shields around the outer reactor surface for reducing radiation heat losses. Reaction rates ωn in eq 14 depend on the surface concentrations of the species and the temperature of the catalyst. To determine these concentrations for each knot of the difference grid, we used the stabilization method to solve the following system of nonlinear algebraic equations s ) ) 2ω1mO2 βSVF(xO2 - xO 2
6. Mathematical Model To describe and numerically investigate the processes occurring in the radial reactor, we considered a onedimensional model with the following assumptions: (i) The radial reactor has a cylindrical symmetry. (ii) The reaction of partial methane oxidation proceeds through the stages of complete methane oxidation (reaction 2), steam reforming (reaction 3), and the water-gas shift (reaction 4). All reaction steps occur independently and simultaneously. (iii) The internal mass transfer control of reaction rates in the catalyst strips is negligible. (iv) External mass transfer control is taken into account, and the effect of multicomponent diffusion is insignificant (because the gas mixture contains 50-60% nitrogen). (v) Plug flow is assumed with respect to the gas phase, and heat is transferred along the catalyst as a result of heat conductivity. (vi Convective heat transfer between the inner wall of the gas-distribution tube and the gas-air flow is not significant. Therefore, the temperature of the inlet gas mixture in the catalyst bed is equal to the temperature of the inlet gas in the gasdistribution tube. Below are given the equations for the material and heat balances:
t g 0, R1 e r e R2 F
∂xi ∂xi +G ) βSVF(xsi - xi) ∂t ∂r
(12)
where i ) {O2, CH4, CO2, H2O, CO, H2}
cgF (1 - )csFs
∂Ts ∂t
1 ∂
( )
r ∂r
rλs
∂Ts ∂r
βSVF(xCO - xsCO) ) (-ω2 + ω3)mCO s βSVF(xCO2 - xCO ) ) (-ω1 - ω3)mCO2 2 s βSVF(xH2 - xH ) ) (-3ω2 - ω3)mH2 2
To describe the stage of complete methane oxidation (reaction 2), we used the kinetic model for a 0.4% Pt/ Al2O3 catalyst,21 adjusted for a Ni/R-Al2O3 catalyst.6 For the catalyst 12.5%(7% Ni/Al2O3) + 85.5% Ni + 2% Cr, the reaction rate expression was modified by introducing a multiplier Ccat and taking into account the amount of the active component in the structured catalyst bed. Because the catalysts developed in this work differs in composition and activity from the catalysts used in refs 21-23, it is wise to introduce an adjustable parameter, C1, that will be determined from the experimental results. Therefore
(13)
[
s s k1aPCH PO 4 2
s s PO k1bPCH 4 2 s s 1 + KCH4PCH + KO2PO 4 2
xi ) x0i , Tg ) T0g
The boundary conditions for eq 14 are
(
(14)
)
-86 000 RTs
k1a ) 8.11 × 105 exp
3
The inlet conditions for eqs 12 and 13 are
+
s s 2 (1 + KCH4PCH + KO2PO ) 4 2
+ RSV(Tg - Ts) -
∑ ωn∆Hn n)1 r ) R1:
s βSVF(xH2O - xH ) ) (-2ω1 + ω2 + ω3)mH2O 2O
ω1 )
∂Tg ∂Tg + cgG ) RSV(Ts - Tg) ∂t ∂r )
s βSVF(xCH4 - xCH ) ) (ω1 + ω2)mCH4 4
(
)
27 300 RTs
KCH4 ) 0.126 exp
(
k1b ) 6.82 × 105 exp
)
-86 000 RTs
]
CcatC1
Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004 4727
(
KO2 ) 7.87 × 10-7 exp
)
92 800 RTs
A kinetic expression for the stage of methane steam reforming (reaction 3) was chosen according to ref 22. Because this kinetics was derived for nickel foil, it was adjusted for our catalyst by taking into account the specific nickel surface area, SNi (for GIAP-3), and the content of the catalyst itself, Ccat, per unit volume of the structured catalyst bed. The correction multiplier C2 was used here as well
{
s k2PCH 4
ω2 )
s PH 2O
s PH 2O
+
[
1-
s 3 ) PsCO(PH 2 s s Keq2PCH PH 4 2O
s 2 b2(PH ) 2
+
s 3 b3(PH ) 2
(
8.12 × 103 exp b2 )
]
}
SNiCcatC2
)
81 789 RTs
Ts3
(
1.82 × 107 exp b3 )
)
195 673 RTs
6.5
Ts 3 × 1013 exp
k2 )
(
)
- 141 287 RTs
Ts3
The kinetic expression for the water-gas shift reaction (reaction 4) was taken from ref 23, where it was derived for 8.7% Ni/R-Al2O3. This expression was adapted for our catalyst by taking into account the content of the active component Ccat per unit volume of the structured catalyst bed
(
ω3 ) k3PsCO 1 -
s PH 2
s PCO 2
s Keq3PsCO PH 2O
(
k3 ) 245 exp
)
)
Ccat
-54 500 RTs
For the chosen experimental conditions and model parameters, the preliminary calculated value of the water-gas shift reaction constant, k3, is very large, so the reaction quickly reaches the quasi-equilibrium state (ω3 being slightly negative at the entrance of the bed). For this reason, in the following calculations, we assumed that ω3 ) 0 and used the resulting equilibrium as an algebraic closure law for the system of differential equations 12-16. The equilibrium constants for steam methane reforming (reaction 3) and the water-gas shift (reaction 4) were taken from ref 15
ln Keq2 ) 8.752 ln Ts -
22 635.63 - 29.768 Ts
5.269 261 × 10-3Ts + 4.927 824 × 10-7Ts2 + 8.736 × 10-12Ts3
ln Keq3 ) -0.768 535 ln Ts +
4943.27 - 1.5062 + Ts
3.010 18 × 10-3Ts - 9.6605 × 10-7Ts2 + 1.475 × 10-10Ts3 To find a steady-state solution, a set of partial differential equations (eqs 12-14) was numerically solved by the stabilization method. In particular, the parabolic eq 14 with the boundary conditions in eqs 15 and 16 was solved using a three-diagonal “chaser” method. The iterative procedure was employed to obtain the joint solution of eqs 12 and 13 with eq 14. The radial heat conductivity coefficient was estimated from the correlation in eq 7 and Figure 3. For the calculation of heat- and mass-transfer coefficients, we used the experimental data presented in Figure 4 and the empirical correlations in eqs 9 and 10. In this case, the length of channels, Lc, was 10 mm, and the hydraulic diameter, dh, was 3 mm. Therefore, β ) DSh/dh, where
[
]
dh Sh ) 1.5 Re Sc Lc R)
(
0.33
λgNu dh
)
dh Nu ) 1.5 Re Pr Lc
0.33
In calculating the effective Reynolds number, we took into account the increase of the gas velocity due to the tortuosity (τt) and porosity ( ) 0.5) of the structured catalyst bed
τt Re ) Ref ,
Ref )
Gdh µ
The value of tortuosity for the developed structured catalyst geometry was approximately defined to be τt ) 10. To calculate the diffusion coefficients and physical properties of the gas-air mixture, we used the database presented in ref 24. 7. Numerical Analysis of the Mathematical Model The main goals of the numerical experiments were (i) adjustment of the parameters of the mathematical model and description of the experimental data, (ii) analysis of parametric sensitivity, and (iii) study of the processes occurring in the radial reactor using a mathematical model. The following figures present the distributions of temperature and species concentrations across the catalyst layer that were calculated using a mathematical model and the experimental data obtained for inlet flow rates of methane and air of 300 and 1292 cm3/s, respectively. This corresponds to a methane concentration of 19.3 vol % in the inlet mixture with air and η ) 0.44. Figure 7 shows the calculated stationary profiles of the gas and catalyst temperatures and the experimental temperatures at the catalyst bed inlet and outlet. As follows from the figure, the experimental and calculated temperatures for the catalyst are quite close. A rather good agreement between the calculated and
4728 Ind. Eng. Chem. Res., Vol. 43, No. 16, 2004
Figure 7. Comparison of the gas and catalyst temperatures calculated with a mathematical model (solid lines) and experimental data (points). Fraction of methane in the inlet air-gas mixture ) 19.3 vol %, coefficient of air excess η ) 0.45.
Figure 8. Comparison of the dry molar fractions calculation for gas mixture species (solid lines) with the experimental concentration measured at the catalyst layer outlet (points). Conditions: natural gas flow rate ) 300 ncm3/s, air flow rate ) 1292 ncm3/s, fraction of methane in the inlet methane-air mixture ) 19.3 vol %, coefficient of air excess η ) 0.454.
experimental inlet catalyst temperatures was obtained by adjusting the multiplier values to C1 ) 3 and C2 ) 0.4 for the rate constants of reaction steps 1 and 2. Significant temperature differences between the gas and solid phases at the inlet catalyst layer are caused by the strong outer diffusion limitation on the rate of methane oxidation. At the inlet, the excess of catalyst temperature over the gas temperature is maximal and equals 1000 °C, which suggests a considerable effect of exothermic reaction 2. As follows from the figure, the temperature profile decreases along the catalyst layer, which is associated with the occurrence of methane steam reforming (reaction 3). The temperature difference in the catalyst layer is ca. 200 °C. Figure 8 shows the profiles calculated for the molar fractions of the gas mixture species and the dry molar fractions measured at the outlet. Note that the experimental and calculated data on CH4, CO, CO2, and H2 are in good agreement. The outlet mixture composition is close to the thermodynamic equilibrium composition for the corresponding outlet temperature. The profiles of concentrations of methane and oxygen are almost the same from a mathematical standpoint, which is associ-
Figure 9. Calculated profiles of the molar fractions of methane and oxygen in the gas-air flow and of methane, oxygen, water, and CO2 on the catalyst surface. Conditions: natural gas flow rate ) 300 ncm3/s, air flow rate 815 ncm3/s, methane fraction in the initial mixture ) 26.6 vol %, coefficient of air excess η ) 0.285.
Figure 10. Calculated profiles of the apparent molar rate of reactions: (1) complete oxidation (2) and steam reforming. Conditions: natural gas flow rate ) 300 ncm3/s, air flow rate ) 1292 ncm3/s, methane fraction in the initial mixture ) 19.3 vol %, coefficient of air excess η ) 0.454.
ated, as shown below, with the almost identical rates of the proposed first two steps of the partial methane oxidation mechanism (i.e., deep methane oxidation and steam methane reforming). Figure 9 shows the calculated profiles of the molar fractions of methane and oxygen in the gas flow and on the catalyst surface. A significant difference in the concentration of each species in the flow core and on the surface suggests a strong effect of the interphase mass transfer on the reaction rate. Moreover, at the catalyst bed inlet, both reactions (methane oxidation and steam reforming) occur under conditions of strong external mass transfer control of their rate by oxygen and methane. It should be noted that the external transport limitation provides a rather high steam/ methane ratio at the catalyst surface (Figure 9), even though there is no steam in the inlet feed. This might be the reason that the process proceeds without coking. This suggestion is of great importance in understanding the effect of transfer processes on the rate of partial methane oxidation. Figure 10 shows the calculated profiles of the rates of the proposed steps of the reaction of partial methane oxidation: (1) complete oxidation and (2) steam conversion. The profiles of the first two steps are quite similar
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in temperature along the catalyst bed (Figure 11) and an increase in the fraction of unconsumed methane and the concentration of hydrogen at the synthesis gas generator outlet (see Figure 12). This conclusion agrees well with the results reported in Tables 2 and 3. The fact that the experimental outlet temperature differs from the calculated data by 76-88 °C (for versions 2 and 3) can be attributed to the effect of heat losses from the outer surface of the structured bed, which seems too difficult to take into account in the boundary condition. The effect of heat losses is pronounced only for low air flow rates (low values of η) or low power releases, because of the increasing relative impact of heat losses in this case. 9. Areas of Application Figure 11. Impact of the air flow rates of 1292, 958, and 815 ncm3/s (versions 1, 2, and 3, respectively) on the temperature distribution in the radial catalyst bed. Points, experiment; continuous line, modeling results. The natural gas flow rate is 300 ncm3/s. Numbers correspond to air flow rates and air excess values as follows: (1) 1292 ncm3/s, η ) 0.454; (2) 958 ncm3/s, η ) 0.33; (3) 815 nm3/s, η ) 0.285.
Figure 12. Impact of the air flow rate and air excess on the radial temperature distribution of methane and hydrogen. The natural gas flow rate is 300 ncm3/s. Numbers correspond variation of the air flow rates as follows: (1) 1292 ncm3/s, η ) 0.454; (2) 958 ncm3/ s, η ) 0.33; (3) 815 ncm3/s, η ) 0.285. Solid lines, calculated results; points, experimental data.
and almost coincide at the inlet. We have numerically shown that only such relation between the exothermic step of deep oxidation and the endothermic step of steam conversion provides a correlation between the calculated and experimentally measured temperatures of the catalyst. 8. Analysis of Parametric Sensitivity Parametric sensitivity is an important characteristic that determines the effect of process conditions on the heat regimes and capacity of the synthesis gas generator. The rates of natural gas and air flow as well as the ratio between them are among such parameters in this numerical analysis. Figures 11 and 12 show the effect of decreasing air flow rate at a constant flow rate of methane. This corresponds to decreasing value of η ) 0.44, 0.33, and 0.29 and increasing inlet fraction of methane ) 19.3, 24.1, and 26.6%. A decrease of η results in a decrease
Synthesis gas is traditionally utilized in methanol synthesis, Fischer-Tropsch processes, and the production of hydrogen for fuel cells. Along with the above areas, our synthesis gas generator can be used in the development of ecologically safe transport vehicles. A rapid increase in the number of automobiles results in serious ecological problems in large cities because of the enormous harmful waste. Traditionally, pollutant emissions are reduced by installing expensive catalytic burners, providing after-burning of CO and reduction of NOx in waste gases. However, the application of threecomponent catalysts increases the cost of cars by 3-4% and reduces the efficiency of their engines by 5%.25 In other words, the current methods for reducing toxic exhausts address the consequences of the principal shortcomings of fuel combustion in the engine. It would be more effective to overcome these shortcomings by changing to new principles of use of hydrocarbon fuels in internal combustion engines (ICEs) and creating a new generation of engine based on a lean burn chamber. This approach is supported by the following facts: When running in urban areas, a vehicle ICE operates mainly at low and moderate loads and therefore emits considerable amounts of toxic combustion products. For example, the average output power of an automobile with an engine rating of 50-100 kW (performance factor of 30%) operating on city roads reaches only 10 kW (performance factor of
