Ind. Eng. Chem. Res. 1995,34, 140-147
140
Conversion of Methanol to Gasoline-Range Hydrocarbons in a ZSM-5 Coated Monolithic Reactor Jimmy E. Antia and m e s h Govind* Department of Chemical Engineering, Mail Location 171, University of Cincinnati, Cincinnati, Ohio 45221
Novel reactor configurations featuring catalysts supported on monolithic or honeycomb structures are being increasingly used for a number of applications. In this work, a zeolite-coated monolithic reactor is employed for the conversion of methanol to gasoline-range hydrocarbons. Experimental results show that the conversion and hydrocarbon product distribution compare favorably with data reported for fxed and fluid beds. Mathematical modeling shows that the conversion here is controlled by diffusion in the molecule-sized intracrystalline pores of the zeolite structure. This finding is of considerable importance because it demonstrates that monolithic reactors are well-suited to zeolite-based catalytic processes.
Introduction The conversion of methanol to gasoline (MTG) over zeolite ZSM-5 catalysts has received considerable attention as an alternate route for the production of transportation fuels (Chang, 1983). The process is highly selective t o hydrocarbons, and the general reaction pathway is represented as follows (Chang and Silvestri, 1977):
2CH,OH
-
+ -
CH,OCH, + H,O C,-C, olefins H,O ~~
paraffins, aromatics, cycloparaffins, C5+olefins This reaction has been investigated in fixed (Chang et al., 1978) and fluidized-bed (Liederman et al., 1978) reactors. Both processes require that the finely synthesized zeolite crystals (1-10 pm) are pelletized for mechanical strength (Satterfield, 1991; Ruthven, 1984). Composite catalyst pellets are formed by molding or extruding mixtures of zeolite crystals and amorphous binders. Larger pellet dimensions are employed in fixed beds to maintain an acceptable pressure drop across the reactor. The exothermic MTG reaction may cause the temperature to rise in fured beds and deactivate the catalyst. This problem is best overcome in fluidized beds, which are operated isothermally. Here, smaller pellet sizes are necessary for easy fluidization. However, fluidized beds report difficulties associated with catalyst attrition and the entrainment of resultant fines. Novel reactor configurations featuring catalysts supported on monoliths or honeycomb structures may overcome the operational problems associated with other reactor types (Irandoust and Andersson, 1988). A monolithic reactor can be imagined as a bundle of small, parallel tubes with catalytic material deposited in the walls. The main advantages of catalyst monolithic reactors are a very low pressure drop across the reactor and a high exposed surface area. Moreover, these reactors can be operated in the horizontal mode without plugging or channeling. Further, the simple flow pattern and reactor structure make for more reliable scale-up. However, for most processes of interest, flow in channels of monolithic reactors is laminar. Monolithic reactors are employed extensively in controlling automobile (Taylor, 1984;Pfefferle and Pfefferle, ~~
* To whom correspondence should be addressed. 0888-5885/95/2634-0140$09.00/0
1987) and industrial NO, emissions (Bartholomew, 1975). Other applications include methanation of carbon oxides (Tucci and Thomson, 1979;Jarvi et al., 19801, steam cracking of naphtha (Heynderickx et al., 19911, and control of volatile organics on immobilized biofilms (Utgikar, 1993). It should be noted that most monolithic reactor applications listed above involve surface catalysis. Under laminar flow, mass transfer to the catalyst-coated walls limits conversion. By contrast, in zeolites most catalytic sites are thought to be located at the intersections of the molecule-sized pores inside the crystal (Derouane and Vedrine, 1980;Jacobs and von Ballmoos, 1982). Here, diffusion of reactants and products into and out of the intracrystalline pores is essential to the process. Configurational or intracrystalline diffusion is orders of magnitude smaller than bulk or Knudsen diffusion, and exists in all zeolite processes regardless of the reactor type. Hence, conversion in conventional pelletized reactor systems may be constrained by configurational and, in the event of very fine intercrystalline spaces, Knudsen diffusional resistances. By comparison, a thin layer of pure zeolite on monoliths offers shorter diffusion distances. Traditional methods of coating zeolites on monoliths require the use of binders to attach the catalyst crystals on the surface (Suzuki, 1989). Higher zeolite loading requires use of larger amounts of binder, which may induce additional intercrystalline mass transfer resistances (Patil and Lachman, 1988). Recently, a method of forming zeolites in situ on the monolith surfaces has been reported (Lachman and Patil, 1989). This ensures that the monolith surface is covered only with the active catalyst. The potential of these reactors, however, has not been adequately studied. In this paper, the application of zeolite ZSM-5 coated monolithic reactors for the production of gasoline-range hydrocarbons from methanol is investigated. The effects of temperature, methanol flow rate, and feed dilution on conversion are studied. Experimental results obtained here are compared with literature data reported for fixed and fluid bed reactors. Mathematical modeling provides insight on the nature of the limiting mass transfer resistance in these reactors.
Experimental Section Catalysts. Cordierite, the ceramic honeycomb used here as catalyst support, was obtained from the Corning
0 1995 American Chemical Society
I
1
I
I
I
I
I
1
1
I I' i , JJL
Table 1. Characteristics of Zeolite Coatings on Monoliths monolith 1 monolith 2 15.63 15.47 wt of zeolite coating (g) 3.73 4.70 0.0158 av uncoated wall thickness (cm) 0.0158 0.0034 0.0043 av coating thickness (cm) 5.3 av crystal length @m) 6.1 0.1014 av coated monolithic channel diam (cm) 0.1032 24.3 surface sioz/&o3 (by EPMA) 35.3 37.1 porosity of washcoat (%) 35.8 w t of uncoated monolith (g)
Glass Works. The uncoated monoliths have 62 channels/cm2, a length of 7.7 cm, and a diameter of 2.5 cm. The zeolite layer was formed on the cordierite surface by a method similar to the one described by Lachman and Patil(1989). The ZSM-5 coating was converted to the hydrogen form via ammonium ion exchange followed by calcination at 600 "C. Two HZSM-5 coated monolith samples were fabricated for this work, and their relevant characteristics are listed in Table 1. Powder X-ray (Cu Ka radiation) diffraction patterns were obtained by crushing the coated monoliths after completing experiments. The diffraction patterns for the coated monoliths are observed to be a superimposition of the individual patterns for zeolite and cordierite (Figure 1). The zeolite pattern best matches that reported for pentasil "silicaliten-an aluminum-free form of ZSM-5. Scanning electron micrographs (SEMs) were employed for "rough" characterization of the zeolite coatings before MTG experiments. SEMs of the uncoated honeycomb as well as the two coated monoliths are presented in Figure 2. Interpenetrating ZSM-5 crystals are observed in the coatings, which comprise spherulitic particles on the first monolith and euhedral particles on the second. ZSM-5 morphologies similar to those obtained in this work have been reported in the literature (Jacobs and Martens, 1987; Satterfield, 1991). Feed. Methanol obtained from Fisher Scientific (HPLC grade) was used as feedstock without further
A
Cordierite A
processing. In one set of experiments, pure methanol was employed. In the other, methanol was diluted with nitrogen. Apparatus and Procedure. A schematic diagram of the laboratory-scale setup, in which experiments were conducted, is presented in Figure 3. Before introduction of the methanol, the reactor was purged with nitrogen for at least 1 h. After the nitrogen flow was shut off, methanol feed was introduced to the preheater by a syringe pump at two rates: 0.075 and 0.15 mumin. All experiments reported here were conducted at 101.3 kPa (1atm) and 300-425 "C. Temperature was measured by thermocouples located at the inlet and exit of the reactor. To ensure steady state, product samples were withdrawn 1 h after the methanol feed was introduced. Reactant and product analyses were conducted in a Hewlett Packard gas chromatograph (5890)-mass spectroscope (5970) system using ultrapure helium as carrier gas. A 150 m long capillary column (Petrocol DH150, Supelco) was employed to separate the product components. The column temperature was initially maintained at 35 "C for 125 min, then heated to 200 "C at 2 "C/min, and maintained at 200 "C for 10 min. Catalyst deactivation, by coke formation, was observed after 4 h of continuous operation. Most of the catalytic activity was restored by burning the coke in a stream of air (16-20 mumin) a t 430 "C for 10-12 h. Hence, data reported in this work should be considered as fresh catalyst activity.
Mathematical Model The complexity of the MTG reaction and the many products formed makes detailed kinetic analysis difficult. Hence, no attempt is made to predict the hydrocarbon product distribution and only reactant conversion is evaluated here. Kinetics. Many reduced models have been proposed (Chang, 1983),where groups of similar compounds were
142 Ind. Eng. Chem. Res., Vol. 34,No. 1, 1995
The activation energy for the first step of the MTG reaction is reported to be approximately 20 kcdmol (Doelle et al., 1981;Chang, 1983). Mass Transfer. Three regimes of mass transfer resistances control of the flow of gases to and from the catalytic active sites: First, there is bulk resistance in monolithic channels as reactant diffuses to the surface of the zeolite coating through the laminar film. This is followed by diffusional resistance to gas flow between the individual zeolite crystals. Finally, the gases diffuse through the molecule-sized pores inside the zeolite crystal to the active sites a t the pore intersections. The reaction products return to the bulk phase in the reverse sequence. On the basis of micrograph measurements, the minimum intercrystalline distance observed in this work is 0.5 pm. At these sizes, diffision of gases between crystals is in the bulk regime. Even under conditions of severe tortuosity in the intercrystalline region, gas flow here would not be hindered as strongly as inside the zeolite crystals. At process conditions, the effective diffisivity in the intercrystalline region is calculated to be 2 orders of magnitude greater than intracrystalline diffusivity. Also, other researchers have demonstrated that for crystals larger than 1 pm, 99% of the active sites are located inside the zeolite (Gilson and Derouane, 1984). Hence, the mass transfer resistance acting in the zeolite coating on the monolithic substrate is probably controlled by intracrystalline or configurational diffusion. Configurational diffusivity is an activated process and is highly temperature dependent (Ruthven, 1984):
I I
I
The activation energy used here is 8 kcdmol, which is the reported value for methanol and, the main product, water (Vigne-Maeder and Auroux, 1990). The mass transfer coefficient for laminar flow of bulk gas in square-shaped monolithic channels is estimated by the following equation (Hawthorn, 1974;Irandoust and Andersson, 1988):
(3)
Figure 2. (a, top) Micrograph of uncoated monolith. (Magnification: 2240x. Reproduced a t 70% of original size.) (b, middle) Micrograph of zeolite coating on monolith 1. (Magnification: 2260x. Reproduced at 70% of original size.) (c, bottom) Micrograph of zeolite coating on monolith 2. (Magnification: 2270x. Reproduced at 70% of original size.)
lumped together into a single species. Disappearance of the “oxygenate” species (comprising methanol and dimethyl ether) followed first-order kinetics for the MTG reaction path (Chang, 1983;Chang, 1984): oxygenates
k
olefins
k2
paraffins
+ aromatics
Temperature dependence of the kinetic constant is of the Arrhenius type:
Reactor Model. One-dimensional mathematical models have been successfully used to explain the behavior of monolithic reactors for automobile catalytic converters (Hegedus, 1975;Heck et al., 1976;Bensalem and Ernst, 1982), methanation of carbon monoxide (Jarvi et al., 1980),and selective catalytic reduction of nitrogen oxides (Buzanowski and Yang, 1990;Beeckman and Hegedus, 1991). A mathematical model is formulated below for the conversion of methanol across a single monolith channel. The model accounts for diffusion and reaction in the zeolite coating, as also possible axial dispersion in the reactor. At any fixed axial position along the reactor, the mass balance for methanol (denoted by A) over a zeolitecoated monolith substrate of thickness dx is represented as follows:
Ind. Eng. Chem. Res., Vol. 34, No. 1, 1995 143
\
\
%lite ?oat& Monolith
Reactor Construction Schematic
AIR
Flowmeter
\
\
-1
I
Syringe Pump
/
Re'actor
Furnace
-
Figure 3. Schematic diagram of experimental setup.
(4)
with the boundary condition at the coating surface (i.e.,
x = 0) For any type of boundary conditions, eq 9 can be solved as outlined by Wehner and Wilhelm (1956): and at the half-wall thickness (i.e., x = h )
It should be noted that the diffusivity, D, is a combined inter4ntracrystalline resistance in the zeolite layer. Equations 4-6 are solved to obtain the concentration at the surface of the zeolite coating:
where
ri
(7)
where Thiele modulus @ = h
8
The mass balance in the axial direction reactor may be written as follows:
(2) of
the
Letting 5 = z/L, z = L/u, and substituting for CA,from eq 7, eq 8 becomes
It should be noted that in the absence of axial dispersion, eq 10 reduces to the form presented by Buzanowski and Yang (1990). The main assumptions made here are listed as follows: 1. Feed and product vapors behave like ideal, incompressible gases. 2. Volume flow of gas is equal in all monolith channels. 3. The pressure drop across the reactor, calculated by the Hagen-Poiseuille equation, is very low ( < 1 mm water). 4. The flow of gases in the monolith channels is laminar. 'Reynold's number = dhu@/,= 0.16-0.43. 5. The flow is fully developed in all but a very small section of the reactor at the channel inlet ( 99%). A n indication that a mass transfer regime is encountered at high temperatures ('400 "C) is favored by comparisons of experimental with predicted conversion values in Table 4. Similar reactor behavior has been reported by investigators for other systems (Satterfield, 1991). This view is further supported by evaluating the change in the Thiele modulus per characteristic catalyst
Ind. Eng. Chem. Res., Vol. 34,No. 1, 1995 145
Methanol feed 0.075 0.15ml/min ml / min
-I
Observed
........... paraffins --_--_olefins .__._.. aromatics
Predicted
o
I_..._..
.I.,.
. . . . . . . . . . . . . . . . . . . . . . .
50
31 0
330
370
350
310 320 330 3 4 0 350 360 370 3 8 0 390 4 0 0 4 1 0 4 2 0 430
410
390
Temperature (C)
Temperature (C)
i270: * 60 :---- ... L
V
\,,,//'
--\
.
j
a
40:
-.'-.
--.-..-. _.-
-2^.+.
D u 50.
1
..-........... ............... ............. " ";'..._........... x) -.,-..._--...........
P 30-
...........,A:...
'.
---/
c_------
...........
Methanol feed Observed 0.075 ml / min o 0.15ml/min o
Predicted
300 310 320 330 340 350 360 370 380 390 400 410 420 430
Temperature (C) Figure 6. Effect of temperature on methanol conversion in (a, top) monolith 1 and (b, bottom) monolith 2 at two methanol feed rates. 7
- , 200
i
250
300
350
400
450
500
550
600
650
Temperature (C) Figure 6. Effect of temperature on Thiele modulus per characteristic catalyst length. Table 4. Conversion at Temperatures > 400 "C at Three Feed Rates methanol flow rate (mumin) 0.075 0.15 0.30
methanol conversion (%) monolith 1 monolith 2 exptl pred exptl pred 99.41 97.36 89.57
99.79 98.23 89.77
99.25 96.46 89.52
99.81 98.33 90.09
length (c$/h)with temperature (Figure 6). At lower temperatures, chemical kinetics predominate, while at higher temperatures, configurational diffusion is dominant. The optimal process temperature, selected at the onset of mass transfer control, is about 370 "C. This compares favorably with literature values (Chen et al., 1989). The effect of temperature on product distribution is plotted in Figure 7. As the temperature is raised,
------
lo:
2 ' 0 1
........ .
,
.
,
.
.-
-ec
,
.
,
.
,
paraffins olefins aromatics
.
,
Figure 7. Effect of temperature on hydrocarbon product distribution in (a, top) monolith 1and (b, bottom) monolith 2. Methanol feed = 0.075 mumin.
increases in light olefins and methane are observed as a result of secondary cracking reactions. Higher temperatures are also reported to uncouple olefin formation reactions from aromatization reactions (Chang et al., 1984). Hydrocarbon Selectivity. The overall conversion and hydrocarbon product distribution obtained in the ZSM-5 coated monolithic reactors at 380 "C and atmospheric pressure are presented in Table 5. Results reported for fixed and fluid bed reactors, operated under similar conditions, are also included in this table. Oxygenate conversion is uniformly high in all three reactor types (299%). Further, it is seen that the hydrocarbon product distribution in monolithic reactors is between that observed for fixed and fluid beds. This may be explained in part by the interaction of differing mass transfer resistances in the three reactor types. In fixed and fluid beds, the gas flow outside the pellets is normally turbulent and, hence, bulk diffusion resistance is rarely present. However, much of the zeolite is inside the pellet and access to the crystals is through pellet pores whose dimensions may be the same as or smaller than the mean free path of gaseous reactant/ product molecules (Palekar and Rajadhyaksha, 1986). Hence, the combined effect of inter- and intracrystalline diffusional resistances in catalyst pellets affects the distribution of hydrocarbon products. Larger catalyst pellets employed in fxed beds result in longer diffusion paths. This might explain the lower amount of intermediate olefins and higher quantity of aromatics in fixed bed reactors compared to fluid beds. For zeolite monolithic systems in this work, only configurational diffusional resistances exist. Like fluidized beds, intercrystalline resistances here are lower than in fxed beds. Under similar process conditions, the product selectivities in monolithic and fluid beds are comparable: monolith 1 and fluid bed in Table 4.
146 Ind. Eng. Chem. Res., Vol. 34,No. 1, 1995 Table 5. Comparison of Methanol Conversion to Hydrocarbons in Different Reactors (P = 1 atm) fured bed (Chang et al., 1978)
monolith 1 (this work)
reactor type monolith 2 (this work)
fluid bed (Liederman et al., 1978)
temperature ("C) WHSV (h-l) SUAI (moVmo1) conversion (%)
371 0.8 -15.0 991
380 0.95= 12.2 98.93
380 0.64= 17.7 99.08
371 1.0 -15.0 99.7
hydrocarbon distrib (wt %) methane ethane propane n-butane isobutane C Z - C ~olefins Cg+ nonaromatics aromatics total
4.1 13.7 4.1 16.8 2.6 14.3 44.4 100.0
0.89 4.91 1.69 6.55 15.29 31.64 39.03 100.00
1.40 7.49 2.67 8.49 12.27 13.94 53.74 100.00
1.03 2.98 1.28 12.98 17.90 37.17 26.66 100.00
% durene in HCs
0.8
4.10
3.61
3.03
+
Based on weight of zeolite.
.....*.-..
..--.+904 0
reactors. No problems of catalyst attrition or excessive pressure drops are observed. 4. At the crystal sizes synthesized (5-7 pm), configurational diffusion plays an important role in determining conversion levels. There have been suggestions that, at submicron particle sizes, the number of catalytic sites on the surface of the zeolite crystal becomes significant (Chen et al., 1989). However, at these reduced sizes, surface reactions may dominate and the unique "shape-selective" aspect of zeolites may be considerably diminished.
Monolith 1 Monolith 2
. . . . . . . . . . . . . . . . . . . . 50
100
150
200
250
300
350
400
450
500
Time on Stream (min) Figure 8. Catalyst aging with time at 370 "C. Methanol feed = 0.075 mUmin.
However, unlike fluid beds, flow past the catalyst surface is laminar in monoliths. Thus, products are transferred to the bulk gas stream by diffusion-a slower process than convective transport. This may account for the lower olefin intermediates observed in monolith 1. In monolith 2, operated a t a lower space velocity, higher quantities of paraffins and aromatics (corresponding to lesser amounts of olefins) were formed. This is probably due to longer residence times in the reactor, whereby the olefins formed undergo further cracking and hydrogen transfer (Chen et al., 1989). Aging Studies. Oxygenate content in the product provided a convenient method of measuring catalytic activity. Aging experiments were conducted at 380 "C with a methanol feed rate of 0.075 mumin. Changes in oxygenate conversion with time are plotted in Figure 8. The catalytic activity is observed to be unchanged for the first 4 h of continuous operation.
Conclusions Important findings of this work are listed as follows: 1. Zeolite coatings of high purity and crystallinity can be consistently formed on ceramic monolithic substrates. Catalyst loadings up to 31% by weight of uncoated monoliths are obtained. 2. Methanol is converted to gasoline-range hydrocarbons with yields similar to those obtained in fixed and fluid bed reactors. The product distribution, though, is somewhat different from standard reactor types. 3. The unique, one-piece construction of monolithic reactors offer advantages of both fixed and fluid bed
Nomenclature a = defined in eq 11 of text CA = concentration of reactant A (mol/cm3) C h = concentration of reactant A in bulk (mol/cm3) C ~=O initial concentration of reactant A in bulk (mol/cm3)
CA,= concentration of reactant A on surface of zeolite (mol/ cm3) dh = hydraulic diameter of the coated monolith channel
(cm) D = combined inter4ntracrystalline diffusivity in the zeolite coating (cm2/s) Db = bulk gas diffusivity (cm2/s) D, = configurational diffusivity (cm2/s) E, = configurational activation energy (cal/mol) E, = kinetic reaction activation energy (cal/mol) h = half-thickness of wall (=characteristic length) (cm) k , = mass transfer coefficient for monolith reactor (cm/s) k = kinetic rate constant (cm3/(gof cat-s)) ko = preexponential kinetic factor (cm3/(gof cat-s)) L = length of reactor (cm) R = universal gas constant ( d . 9 8 7 cal/(mol-K)) T = reaction temperature (K) u = linear velocity of gas in the monolith channel (cm/s) X A = conversion of oxygenates x = radial position in the zeolite layer (cm) z = axial position along reactor length (cm) 5 = dimensionless axial length 6 = density of zeolite catalyst (g/cm3) e = density of gas (g/cm3) p = viscosity of gas (g4cm.s)) r = reactor residence time ( 8 ) u = area per unit volume of reactor (cm2/cm3) $J = Thiele modulus as defined in eq 7
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Abstract published in Advance A C S Abstracts, December
1, 1994.
