Znd. Eng. Chem. Res. 1996,34,1581-1587
1681
Oxidative Coupling of Methane in a Vibrofluidized Bed at Low Fluidizing Velocities Africa Santos, Jes6s Santamaria, and Miguel Menhdez* Department of Chemical and Environmental Engineering, Faculty of Science, University of Zaragoza, 50009 Zaragoza, Spain
A fluidized bed reactor has been used to study the influence of the operation parameters on the reactor behavior during oxidative coupling of methane. A vibration device was employed to avoid the agglomeration of the catalyst particles, thus allowing operation at low gas velocities, near the minimum fluidization velocity. An improved selectivity for hydrocarbons was obtained in this way, compared to the results a t high gas velocities.
Introduction Numerous research groups have been involved during the last decade in the effort to obtain more valuable and more easily transportable hydrocarbons from methane. Among the many different possible processes, oxidative coupling of methane has been considered as a promising one (Matherne and Culp, 1992; Renesme et al., 1992; Parkyns et al., 19931, and the development of suitable catalysts has been the subject of extensive research. A review by Maitra (1993) reports several hundred catalysts selected from the copious literature in this field. However, comparatively little effort has been made to develop or test reactors suitable for this process, and with a few exceptions, only fmed bed reactors have been employed. Since methane oxidative coupling is very exothermic, if a fmed bed reactor is t o be used for the process, a multitubular reactor would be required. This would obviously be an expensive solution given the high temperatures of reaction, about 800 "C, and the need to use special materials, not only to accommodate the severe reaction conditions but also to avoid the combustion of methane, which is catalyzed by steel at the reaction conditions. The use of a fluidized bed reactor offers significant advantages over the fixed bed reactor. (i) The first is an easier temperature control, given the isothermicity of the reactor and the high heat transfer coefficients which are typical of fluidized bed operation. (ii) The second advantage is a lower requirement of feed preheating since at least part of the heat necessary to raise the feedstream temperature to the reaction temperature can be provided in a fast and efficient manner as it enters the reactor. In this way, an important saving in heat exchanging equipment may be obtained. (iii) At least in some cases, as it has been found experimentally (Follmer et al., 1988;Andorf and Baerns, 1990) and by modeling studies (Mleczko et al., 1991; Menendez and Aurensanz, 19931, the fluidized bed reactor can provide a higher selectivity for hydrocarbons than the fxed bed reactor. Adris et al. (1991) speculated about the possibility of an improved performance of fluidized bed reactors with respect to fixed bed reactors in the case of equilibriumlimited kinetics because of the membrane effect caused by the bubbles. In this case, the bubbles would remove part of the reaction products from contact with the
* To whom correspondence should be addressed. E-mail:
[email protected]. Fax: (+34)(76)567920.
catalyst in the emulsion phase, therefore driving the forward reaction beyond equilibrium limitations. In the case of methane oxidative coupling, a reason for the improvement in selectivity may be the restriction to mass transfer between the bubbles and the emulsion phase (Andorf and Baerns, 1990; Mleczko et al., 1991; Menendez and Aurensanz, 1993). This would result in a low oxygen concentration in the emulsion phase, with the corresponding increase in selectivity, since it seems well-established that low oxygen concentration favors the selectivity for hydrocarbons due to the higher reaction order of oxygen in the combustion reactions with respect to the coupling reactions. In fact, a low oxygen concentration has been at the basis of some improved reactors for oxidative coupling, where the oxygen feed is distributed, either a t various discrete points (e.g. Choudhary et al., 1989; Fino1 et al., 1994) or continuously along the reactor using a porous membrane (Lafarga et al., 1994; Coronas et al., 1994). Also, Rizayev et al. (1994)have found that the selectivity in oxidative dehydrogenation of C4 paraffins is higher in a fluidized bed than in a fixed bed reactor with cofeeding of reactants or even higher than that obtained in a three-stage system with distributed oxygen feed. (iv) Other advantages of fluidized bed reactors are the easy addition and withdrawal of catalyst from the reactor zone during operation and their relatively large volume/reactor wall surface area ratio, important as wall effects are significative if materials different from quartz or sintered alumina are used. It may be concluded that a fluidized bed reactor constitutes a suitable alternative for a large scale process (Parkyns et al., 1993). However, the fluidized bed reactor has some disadvantages, the main one being the bypass of the fluidization gas, due to the bubbles. In oxidative coupling of methane, a second problem arises due to the increase in gas phase volume, since this reaction also takes place in the gas phase, with a low selectivity (Lane and Wolf, 1988). Also, the backmixing of the gas may reduce the selectivity (Kunii and Levenspiel, 1991) since this reaction is, at least partially, a series reaction, with the hydrocarbons being the valuable intermediate product. A procedure to avoid the above-mentioned disadvantages is to operate at low gas velocities, near the minimum fluidization velocity. Unfortunately, most of the best catalysts contain alkali promoters, and their salts are partially fused at the reaction conditions. This causes the defluidization of the bed when low relative gas velocities (ur,defined as the ratio of the gas velocity to the minimum fluidization velocity)are employed. Due
0888-588519512634-1581$09.0OlO 0 1995 American Chemical Society
1582 Ind. Eng. Chem. Res., Vol. 34,No. 5,1995 Movable
To exit, flow measurement and GC analysis
Electrical fumace tic bed Control Thermowell Quartz distributor
Vibration system
Mass-flow controlled CH4/O2 mixture
Figure 1. Scheme of the reactor employed.
to this reason, most of the results of oxidative coupling of methane in fluidized bed reactors reported to date have been obtained with high relative fluidization velocities. Thus, Edwards et al. (1990)reported results obtained in the 5-17 urrange with a Li/MgO catalyst, using a low lithium loading (about 0.4%) to avoid defluidization. Andorf et al. (1991) operated a t a urof 15 with a PbO/Al203 catalyst because of the same reason. These authors also reported results a t low relative velocities (3.2 and 2.1) with CaO and Na2C03/ CaO catalysts diluted with a-&O3, but the yields of hydrocarbons were rather low. The use of a vibrofluidized bed allows for the use of low gas velocities with solids which tend to agglomerate and is used industrially to roast ores or to dry cohesive and sticky solids. Rizayev et al. (1994) carried out experiments on a vanadium-containing catalyst for oxidative dehydrogenation of C4-C5 paraffins in a pilot scale installation using a vibrofluidized bed reactor. In this paper, a vibrofluidized bed has been employed for the oxidative coupling of methane in order to employ relative gas velocities near the minimum fluidization velocity. Even if the application of a vibrofluidized bed were not technically feasible a t an industrial scale in this process, the results presented herein will provide some insight into the reactor operation a t low relative gas velocities.
Experimental Section The reactors employed were made of quartz, with a diameter of 3 cm and a total height of 30 cm, and used quartz frits as gas distributors. The reactor temperature was controlled by a thermocouple inserted in a quartz thermowell obtained by deformation of the reactor wall. In addition, the temperature profile in the reactor could be measured by means of a movable thermocouple inside an axial quartz thermowell. The height of the catalytic bed depends on the amount of catalyst in the bed and the operation conditions, being about 15 cm in most cases. The reactor was partly introduced into an electrically heated furnace in such a way that about 8 cm of the freeboard stayed out of the furnace. A schematic of the reactor employed is shown in Figure 1. All the experiments reported in this work were carried out with methane/oxygen mixtures, without dilution
with inert gases. The reactant gases were mass flowcontrolled, and the set point temperature was controlled within f 5 K. The pressure drop could also be measured both before the distributor plate and inside the bubbling bed. To improve the fluidization behavior, two procedures were tested, namely, the addition of a second solid (ground quartz) and the use of a vibration system at the bottom of the reactor. To induce the vibrations, the inlet quartz tube was coupled to an eccentric rotor by means of a flexible connection. The rotation was controlled to give a frequency of approximately 3 s-l and an amplitude of about 1 cm. By using this device, a good fluidization was obtained, even at gas flow rates near the minimum fluidization velocity. The exit gases were directed to a cold trap and then to an on-line gas chromatograph where they were analyzed for CH4, 02, C02, C2H6, C2H4, and hydrocarbons with three and four carbon atoms. After the cold trap, a supplementary flow measurement was provided to measure total exit flow rates. Carbon mass balance closures were always within f5% and usually within f 3 % . Selectivities are reported as the total number of moles of carbon in a given product species divided by the total number of moles of methane reacted. C2 selectivity refers only to ethane and ethylene in the exit gases, while C2+ also includes C3 and C4 hydrocarbons. A Li/MgO catalyst with a lithium loading of approximately 3% by weight has been prepared by the slurry method. A more detailed description of the preparation of the catalyst has been provided elsewhere (Coronas et al., 1994). This catalyst was selected because of its good characteristics as a coupling catalyst and also because Li/MgO catalysts have been employed in numerous works since it was described by Ito et al. (1985) and may be considered as a reference catalyst for oxidative coupling of methane. The particle size employed was between 100 and 500 pm, depending on the particular experiment. The measured BET surface areas were between 0.2 and 0.8 m2/g,depending on the particular batch and on the extent of use of the catalyst.
Fluidization Behavior The Li/MgO catalyst did not fluidize properly at the reaction conditions employed, i.e. temperatures of about 800 "C and a reaction atmosphere containing methane, oxygen, and the reaction products. Under these conditions, the particles tended to agglomerate, partly due to the fact that most of the possible lithium compounds are liquid, and the catalyst particles become sticky. Three options may be considered to improve the quality of the fluidization. (a) The first is to operate at high gas velocities in such a way that the energy provided by the entering gas is enough to break the aggregates of particles. The disadvantages of this procedure have been pointed out above. (b) The second is to add a second inert solid, which acts as a coadjutant of fluidization. To this end, crushed quartz with a particle size similar to that of the catalyst was used. The bed fluidized correctly with a mixture containing 50% by weight of each of the two solids. However, very low selectivity for hydrocarbons was obtained. Figure 2 shows the data obtained with a relatively wide variation of CH4/02 ratios (from 9 to 2). It can be seen that the selectivities in the presence of ground quartz were roughly half of those attained when only the catalyst was present. This low selectivity may
80
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1583
,
I
,e,; 1 ,
--
p
,
Catalyst-Quartz Catalyst I Vibration system
IO 0
5
10
15
25
20
35
30
40
Methane Conversion (70) Figure 2. Hydrocarbon selectivity vs methane conversion in a fluidized bed of W g O catalyst with and without the use of quartz as a coadjutant for fluidization. Conditions: temperature, 800 "C; total flow rate, 1500 cm3 (STP)/min; particle diameter between 100 and 250 ym; catalyst weight, 50 g.
X CH, 0
8 0
I
c)
m
8
a
compared to those in the presence of ground quartz. Therefore, we decided to adopt the vibrofluidized bed for the rest of the experiments reported in this work. An important parameter of any fluidized bed system is the minimum fluidization velocity. To obtain the minimum velocity of gas necessary to obtain a smooth fluidization, in a plot of pressure drop vs gas velocity, the points in which the pressure drop increases linearly with the gas velocity were fitted by a straight line passing through the origin. The minimum fluidization velocity was calculated as that corresponding to the intersection of this straight line with the line of constant pressure drop. The value of umf obtained by this procedure for the particle size used in most of the following experiments (between 100 and 250 pm) was approximately 6.5 cm/s, which corresponds to a gas flow of 750 cm3(STP)/min. It was observed that the pressure drop in the fluidized bed was slightly larger than the static pressure due to the solid weight. This phenomenon is due to the agglomeration of the particles. The apparent viscosity of a fluidized bed with sticky particles is not negligible, and part of the energy of the fluidization gas is used to break the links between particles. In fact, fluidization with these kinds of particles is only possible when the energy transmited by the fluidizing gas to the particles is higher than the binding force due to the adhesion of the particles (Moseley and O'Brien, 1993).
750
0
16.04
0
21.42
A
25.64
X
27.06
0
36.63
fluidized 725
Reactor Behavior 0
I
I
lb
'
I
I
be due t o some catalytic effect of the quartz particles and also due to the formation of lithium silicates, also with catalytic activity and low selectivity. In addition, this operation procedure increases the catalyst deactivation by lithium loss. It has been observed (Korf et al., 1992; Slagtern et al., 1992) that deactivation in the presence of quartz surfaces is faster than in their absence. Thus, for instance, when a mixture of 90% methane and 10% oxygen was fed into the reactor entrance at 800 "C, the yield diminished to about 60% of the initial value in 7 h of operation but only to 95% in the absence of quartz. (c) The third option is to induce a vibratory movement in the bed. The vibration system previously described allows for good fluidization, even at low gas velocities. In addition to visual inspection of the bed, defluidization may be efficiently detected because significant temperature gradients appear in the catalyst bed. Instead, the vibration system provides a good fluidization that ensures almost isothermal operation, even under relatively high methane conversions, as shown in Figure 3, where the temperature profiles have been represented. It can be seen that the maximum temperature difference between any two positions within the bubbling zone is less than 10 "C. Figure 3 also shows that, after leaving the bubbling zone, the gases undergo a very rapid cooling. This is a useful feature since a fast quench of the exit gases should minimize the gas phase reactions in the freeboard. As shown in Figure 2, the vibration system also enabled us to obtain considerably better selectivities,
The operating conditions, i.e. gas velocity, reaction temperature, amount of catalyst in the bed, and feed composition, have been varied in the system described above. The results obtained yield information about the fluid dynamics and kinetic behavior of the reactor. In all the experiments, a smooth fluidization was obtained. Preliminary experiments, in which the temperature profile in the freeboard was varied by changing the relative positions of the reactor and the furnace, showed little effect from the different temperature profiles. Therefore, it seems reasonable to assume that the quenching of the exit gases is fast enough as to avoid significant reaction in the freeboard. Effect of Reaction Temperature. Figure 4a,b shows the variation of the conversion of methane and oxygen with the temperature of the bubbling zone. It can be observed that a t temperatures above 800 "C, most of the oxygen feed to the reactor was consumed, irrespective of its concentration in the feed gas. This gives rise to a plateau in the curves of methane conversion vs temperature. A plot of selectivity vs methane conversion, shown in Figure 5, allows for a comparison between the results obtained a t different temperatures. It may be seen that the line that goes through the data obtained a t 800 "C lies over the rest of the experimental data. This is probably the consequence of a balance between the favorable effect of higher temperatures on the selectivity of the catalytic reaction that is found in most cases (Amenomiya et al., 1990) and the increase in the nonselective gas phase reactions that also takes place with increased temperatures. In view of this, the rest of the experimental work was carried out at 800 "C. The ratio of ethylene/ethane in the reaction products increased with temperature, varying from about 1 a t 750 "C to 4 at 850 "C. This is due to the production of ethylene from both thermal pyrolysis and oxidative dehydrogenation, which increase with temperature. The
1584 Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995
725
750
775
800
825
Temperature ("C)
850
875
0
p
10
20 30 40 Catalyst Weight (g)
C a2 M
8
750
715
800
825
850
875
50
Temperature ("C)
70
I
I
0
Figure 4. Variation of the methane conversion (a) and the oxygen conversion (b) with the temperature of the bubbling zone a t three different CH4/02 ratios. Other conditions as for Figure 1.
X
1500 1125
0
900
0
6o 0
725
60
Q (cm3/min)
x
2.511
50
10
I
I
20 30 40 Catalyst Weight (g)
50
60
Figure 6. Variation of the methane conversion (a) and the oxygen conversion (b) with the amount of catalyst in the bed at three different gas flow rates. C W 0 2 = 5/1, other conditions as for Figure 1.
h
.-x .->
Table 1. Oxygen Conversion in a Short Bed (ca. 1.5 cm depth), Compared with the Ratio u d u , at Two Values of Partial Pressure of Oxygen in the Feed (bar)
65
Y
-
60
U
8
55
0.5 0.667 0.833
0
0 v)
2 V
2 a
50
G
45
G 0
10 20 30 Methane Conversion (%)
40
Figure 6. Hydrocarbon selectivity vs methane conversion at various temperatures and with a fixed flow rate of 1500 cm3 (STP)/ min. Other conditions as for Figure 1.
ability t o obtain ethylene preferentially is important because of its higher price, as compared to ethane. Influence of the Catalyst Weight. If gas reactions in the freeboard can be neglected, varying the weight of the catalyst bed is roughly equivalent to obtaining the concentration profiles along the bed. Figure 6 shows the variation of the methane and oxygen conversions for different amounts of catalyst in the reactor at three values of the gas flow rate. The conversion of methane increases asymptotically along the bed until all the oxygen fed to the reactor is consumed. Similar figures (not shown) were obtained with other concentrations of oxygen in the feed gas. It can be seen that a fast reaction takes place as the gas enters the bubbling bed, with most of the conversion taking place in the first few centimeters of the reactor.
~
U
Xo, (withpo, = 0.25) 0.503 0.705 0.834
Xo, (withpo, = 0.17) 0.556 0.694 0.813
Figure 6 also shows that for a given weight of catalyst, the conversion of oxygen decreases when the gas flow increases, due to the important bypass of gas produced by the bubbles. This effect is most pronounced when a small amount of catalyst is present in the reactor (about 5 g). At this low value of space time, the oxygen conversion roughly correlates with the fraction of the oxygen fed to the reactor that goes initially to the emulsion phase, estimated as u d u ,as may be seen in Table 1. This suggests a picture of the reactor in which the gas that enters the emulsion phase is quickly converted, while the gas that initially enters in the bubbles reacts more slowly, as it is transferred to the emulsion phase. Influence of the Feed Flow Rate. The effect of the feed flow rate on the methane and oxygen conversions is shown in Figure 7a,b. These data were obtained with a fixed mass of catalyst (50 g) in the bed. As could be expected from the results in Figure 6, an increase in the feed flow rate results in a decrease of the oxygen and methane conversions. A faster decrease with the feed flow rate was observed for higher values of oxygen concentration in the feed. Notwithstanding the effect of mass transfer limitations, if the apparent reaction order with respect to oxygen was one, the conversion
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1585 40
"V
0
0
h
5
75
I
Y
5
o
s 10
o
o
o
o
o
o
o
o
o
4 700
900
1100
1300
1500
1700
3 45 CL
I
20
15
1900
I
75
.-8
! 6
90
E
85
E
6
80
-
65
60
-
55
-
900
1100
1300
1500
1700
Q (cm3(STP)/min)
would be expected to be approximately independent from the oxygen concentration. Therefore, the observed behavior is consistent with the experimental observation on Li/MgO catalysts that the apparent reaction order for oxygen is less than unity for both the reactions of methane combustion and methane oxidative coupling (Mirodatos et al., 1988). Figure 8a shows the plot of hydrocarbon selectivity vs methane conversion, using different oxygen partial pressures and space times. It can be seen that the data corresponding to each of the oxygen partial pressures employed are clustered around approximately the same value of selectivity, in spite of the fact that a wide variation of space time, from 0.062 to 0.025 g min/cm3, was used t o obtain the data. This means that the selectivity for hydrocarbons depends only slightly on spatial time and is mainly related to the concentration of oxygen in the feed. This is shown in Figure 8b as a plot of hydrocarbon selectivity vs the partial pressure of oxygen. Figure 8a also shows that the effect of the methane conversion on the selectivity is a relatively minor one. This may seem surprising since the formation of COXis, in part, a series reaction, and thus, the selectivity usually decreases when the methane conversion is increased. Although this effectively takes place, its influence is smaller than that of the partial pressure of oxygen, which can be linked to the different apparent reaction orders of oxygen in the reactions of methane combustion and methane coupling. In any case, the range of methane conversion explored is relatively small at each of the partial pressures of oxygen employed, and a larger effect can be expected when the methane conversion is varied over a wider range. Particle Size of the Catalyst. Figures 9 and 10
I
25
30
35
40
I
-
45
-
8
8 0
0
I)
0.1
1900
Figure 7. Variation of the methane conversion (a)and the oxygen conversion (b) with gas flow rate and with the CH4/02 ratio. Other conditions as for Figure 1.
I
I
0
700
1
I
50
75
I
Methane Conversion (%)
-
70
*-
I
40
Q (cm3(~~)/min) 100
. .I
--w
o
0.3
0.2
P O2(bar)
0.4
Figure 8. Variations in the hydrocarbon selectivity with the methane conversion (a) and with the partial pressure of oxygen in the feed (b). Other conditions as for Figure 1. 35
E
200-320 pm 300-500pm
h
u
33
L'I
0
C
m
.s
s
21
0
1 t
700
1000
1300 1600 1900 Q (cm3(STP)/min)
2200
Figure 9. Methane conversion vs total flow rate using different particle sizes. CHd02 = 3/1, other conditions as for Figure 1.
show the conversion of methane at various gas flow rates and particle sizes and the plot of selectivity vs methane conversion at various particle sizes over a range of CHdO2 ratios from 9 to 2, respectively. It appears that larger particles provide slightly higher selectivity and higher methane conversion at low feed gas velocities (less than 1200 cm3 (STP)/min) and gave roughly the same conversion at larger gas flow rates. The variation of the particle size influences both the fluid dynamics of the bed and the catalytic properties of the solid. The effect is complex since not only will the minimum fluidization velocity and bubble size change, but also, given the low intrinsic surface area of the catalyst employed, the use of lower particle sizes could lead to significant increases in the total surface area. However, calculations with the particle sizes employed show that the increase in the total surface
1586 Ind. Eng. Chem. Res., Vol. 34,No. 5, 1995 80
-
h
i
E
* .-
.3
> u
70
-
60
-
-
x
gx
o
0
'b
. d
a, a,
m
coupling using a fluidized bed reactor. In this figure, lines 3 and 4 give the upper and lower envelope of the results obtained by these authors using a Li/MgO catalyst (catalyst A), respectively. Catalysts B and C (lines 1 and 2, respectively) are proprietary catalysts developed by CSIRO. The main difference between their results and this work is the operation at low gas velocities, attained in this work by the use of the vibratory device which helps to avoid the agglomeration of the bed. It can be seen that most of our results lie over their previous results with Li/MgO and even reach the data obtained with one of the highly optimized proprietary catalysts (catalyst C).
100-200pm 200-320pm 300-500pm
0
O
d +@
.a
50
7BXo
-
x
R z L
40
I
I
I
o I
I
Conclusions The behavior of a fluidized bed reactor for the oxidative coupling of methane can be explained on the basis of the fluid dynamics of the bed and some general kinetics trends. Smooth, isothermal reactor operation at gas velocities near the minimum fluidization velocity can be obtained by using a vibrofluidized bed. This improves the hydrocarbon yield with respect to that obtained in a fluidized bed operating at higher gas velocities. The possibility of influencing the reactor behavior by varying the gas velocity or the particle size opens an interesting field for optimization of the reactor operation.
100 90
80
70 60 E
f
e
50 40
30
1
20
I
I
I
1
I
I
0
5
10
I5
20
25
-n
&'
1
30 Methane Conversion (%)
\l
Acknowledgment
I
35
40
Figure 11. Comparison of the results obtained in this work with other results in the literature.
area would be of a small magnitude in our case (around 5% at the most). Another likely contributing factor is the increase in the amount of lithium promoter per unit of surface area of the support that takes place when the particle size is increased for a given catalyst mass. This would lead to higher initial surface lithium loadings and also, given the lower surface area available, to lower rates of lithium loss. Although the resulting effects of the above factors on the conversion and selectivity are not easily predictable, some general trends can be established. Thus, a larger particle size means a higher fluidization velocity, and, for a given total flow rate, a larger fraction of gas in the emulsion phase and therefore a larger conversion a t the bed entrance. Due to this high initial reaction rate, the oxygen concentration in the emulsion phase is depleted, and further reaction only takes place as additional oxygen is transferred from the bubbles to the emulsion phase. An increase in the particle size should produce fewer and larger bubbles in the bed, from which oxygen would be transferred more slowly. The result would be lower concentrations in the emulsion phase. This effect, coupled with the lower surface area at higher particle sizes could be responsible for the higher selectivities observed at a given methane conversion.
Comparison with Previous Results The results obtained in the above-explained experiments are compared in Figure 11 with the results obtained in a fluidized bed by Edwards et al. (1990, 1992) using three different catalysts. These are probably the best results obtained in methane oxidative
Financial support of DGICYT, Spain (Project PB 9303111, is gratefully acknowledged.
Literature Cited Adris, A. M.; Elnashaie, S. S. E. H.; Hughes, R. A Fluidized Bed Membrane Reactor for the Steam Reforming of Methane. Can. J , Chem. Eng. 1991,69,1061. Amenomiya, Y.; Birss, V. I.; Goledzinowski, M.; Galuszka, J.; Sanger, A. R. Conversion of methane by oxidative coupling. Catal. Rev.-Sci. Eng. 1990,32 (31,163. Andorf, R.; Baerns, M. Oxidative coupling of methane in fluidizedand packed-fluidized-bed reactors. Catal. Today 1990,6, 445. Andorf, R.; Mleczko, L.; Schweer, D.; Baerns, M. Oxidative coupling of methane in a bubbling fluidized bed reactor. Can. J. Chem. Eng. 1991,69,891. Choudhary, V. R.; Chaudhari, S. T.; Rajput, A. M.; Rane, V. H. Beneficial Effect of oxygen distribution on methane conversion and Cz-Selectivity in oxidative coupling of methane to CZhydrocarbons over lanthanum-promoted magnesium oxide. J . Chem. SOC.,Chem. Commun. 1989,1526. Coronas, J.; Menbndez, M.; Santamaria, J. Methane oxidative coupling using porous ceramic membrane reactor. Part 11: Reaction studies. Chem. Eng. Sci. 1994,49,2015. Edwards, J. W.;Do, K T.; Tyler, R. J. Reaction Engineering studies of methane coupling in fluidized-bed reactors. Catal. Today 1990,6,435. Edwards, J. H.;Tyler, R. J.; White, S. D. Oxidative coupling of methane over lithium-promoted magnesium oxide catalysts in fured-bed and fluidized-bed reactors. Energy Fuels 1990,4,85. Edwards, J. W.;Do, K. T.; Tyler, R. J. The OXCO Process: The Direct conversion of Natural Gas to Olefins and Liquid Fuels Using Fluidized-Bed Technology. In Methane Conversion by oxidative processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; Chapter 13. Finol, C.; Menhdez, M.; Monzbn, A,; Santamaria, J. An experimental study of methane oxidative coupling in fixed bed reactors with a distributed oxygen feed, Chem. Eng. Commun.,in press, 1994. Follmer, G.; Lehmann, L.; Baerns, M. The application of laboratory-scale catalytic k e d and fluidized bed reactors in the oxidative coupling of methane. Am. Chem. SOC.,Diu. Pet. Chem. 1988,33,453.
Ind. Eng. Chem. Res., Vol. 34, No. 5 , 1995 1687 Ito, T.; Wang, J. X.; Lin, C. H.; Lunsford, J. H. Oxidative dimerization of methane over a lithium promoted magnesium oxide catalyst. J. Am. Chem.SOC.1986,107,5062. Korf, S. J.; Roos, J. A.; Ross, J. R. H. The development of doped LUMgO catalyst systems for the low-temperature oxidative coupling of methane. In Methane Conversion by oxidative processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; Chapter 6. Kunii, D.; Levenspiel, 0. Fluidization Engineering, 2nd ed.; Butterworth-Heineman: Boston, 1991; Chapter 12. Lafarga, D.; Santamarfa, J.; Menendez, M. Methane oxidative coupling using porous ceramic membrane reactors. Part I: reactor development. Chem. Eng. Sci. 1994,49,2005. Lane, G. S.;Wolf, E. E. Methane utilization by oxidative coupling. I. A study of reactions in the gas phase during the cofeeding of methane and oxygen. J. Catal. 1988, 113,144. Maitra, A. M. Critical performance evaluation of catalysts and mechanistic implications for oxidative coupling of methane. Appl. Catal. A 1993, 104, 11. Matherne, J . L.; Culp, G. L. Direct conversion of methane to C2's and liquid fuels: process economics. In Methane Conversion by oxidative processes; Wolf, E. E., Ed.; Van Nostrand Reinhold: New York, 1992; Chapter 14. Menhdez, M.; Aurensanz, J. Simulation of a fluidized bed reactor for oxidative coupling of methane. Afinidad 1993, 50, 289. Mirodatos, C.; Martin, G. A. Oxidative coupling of methane on LiMg oxides: Kinetics and mechanism involving carbenic intermediates. Proc.-Int. Congr. Catal., 9th 1988,2, 899.
Mleczko, L.; Andorf, R.; Baerns, M. Simulation of fluidized bed reactors for catalytic oxidative coupling of methane to CZ hydrocarbons. Chem. Eng. Technol. 1991,14, 325. Moseley, J. L.; O'Brien, T. J. A model for agglomeration in a fluidized bed. Chem. Eng. Sci. 1993,48, 3043. Parkyns, N. D.; Warburton, C. I.; Wilson, J. D. Natural gas conversion to liquid fuels and chemicals: where does it stand? Catal. Today 1993, 18, 385. Renesme, G.; Saint-Just, J.; Muller, Y. Transportation fuels and chemicals directly from natural gas: how expensive? Catal. Today 1992, 13, 371. Rizayev, R. G.; Talyshinskii, R. M.; Seifullayeva, J. M.; Guseinova, E. M.; Panteleyeva, Y. A.; Mamedov, E. A. Oxidative dehidrogenation of the C4-C5 paraffins over vanadium-containing oxide catalyst. In New developments in selective oxidation II; Cortes, V., Vic, S., Eds.; Elsevier: Amsterdam, 1994; p 125. Slagtern, A.; Dahl, I. V.; Jens, K. J.; Hansen, L.; Seiersten, M. Reactor materials for use with the Li/MgO methane coupling catalyst. Appl. Catal. A 1992, 91, 13. Received for review July 18, 1994 Revised manuscript received December 23, 1994 Accepted January 31, 1995" IE940447E
* Abstract published in Advance ACS Abstracts, March 15, 1995.
