Oxidative coupling of methane in porous Vycor membrane reactors

j o u r n a l of MEMBRANE SCIENCE

ELSEVIER

Journal of Membrane Science 116 (1996) 253-264

Oxidative coupling of methane in porous Vycor membrane reactors A.M. Ramachandra, Y. Lu, Y.H. Ma, W.R. Moser, A.G. Dixon * Center for Inorganic Membrane Studies, Department of Chemical Engineering, Worcester Polytechnic Institute, Worcester, MA 01609, USA

Received 20 October 1995; revised 16 January 1996; accepted 30 January 1996

Abstract Porous Vycor membrane tubes were used in shell-and-tube type membrane reactors to study the effect on the oxidative coupling of methane of metering the oxygen into the catalyst bed. Experimental studies showed that under conditions of complete oxygen conversion, Vycor membrane reactors packed with Sm203 catalyst exhibited enhanced hydrocarbon (C 2) selectivity. C 2 yields were comparable to those of the conventional co-feed packed bed reactors operated under the same conditions. The higher C 2 selectivity in the membrane reactors indicated that, for methane coupling, regulating the supply of oxygen along the length of the packed bed may be beneficial to C 2 formation. Keywords: Microporous and porous membranes; Inorganic membranes; Glass membranes; Membrane reactors

1. Introduction Conversion of abundantly available natural gas ( C H 4) into valuable products has become an increas-

ingly attractive proposition in recent years with rapid advances in technology. One of the more recent and potentially attractive prospects is the oxidative coupling of methane (OCM) to produce higher hydrocarbons such as ethane and ethylene (C2s). The last decade has seen an exponentially increasing interest in OCM technology as evidenced by the reported work in the literature [1-17]. The OCM process involves a metal-oxide catalyzed reaction of methane in the presence of oxygen to form C 2 hydrocarbons

* Corresponding author. Tel.: (508)-831-5350; Fax: (508)-8315853; E-mail: [email protected].

(C2H 4 and C 2 H 6 ) , along with the oxygenated sideproducts CO, C O 2 and H 2 0 . Increasing methane conversion, however, is accompanied by lowered C 2 selectivity, and the resulting decrease in C 2 yield is characteristic of such reactions where the intermediate products ( C 2 H 4 , C2H 6) are more reactive than the reactants (CH4). An excellent review of the evolution and current state of research in OCM technology is given by Amenomiya et al. [1]. High temperature reactions occurring during methane coupling consist of both a complex set of gas phase homogeneous reactions as well as catalyzed heterogeneous reactions, involving the presence of methyl radicals [1-4]. Various mechanisms of methyl radical activation have been proposed [6,7,9,17], based on the different kinds of oxygen species that could be formed during the high temperature OCM reactions: the lattice oxygen 0 2-, the

0376-7388/96/$15.00 © 1996 Elsevier Science B.V. All rights reserved PII S 0 3 7 6 - 7 3 8 8 ( 9 6 ) 0 0 0 4 4 - 0

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oxygen anion radical (O), the peroxide ion (02-), and atomic oxygen (O). The review by Voskresenskaya et al. [9] summarizes the different views on the roles of these oxygen species in the overall reaction mechanism of methane coupling. Structural defects in the oxide [9], the gas phase oxygen partial pressure [10], the concentration of the deep oxidation products C O / C O 2 [5], as well as other factors all influence the oxidant activation process for methane coupling. For OCM technology to become commercially viable, C 2 yields have to be higher than the presently attainable levels of 20-25%. Although many metaloxide catalysts that exhibit selectivity for the formation of C 2 hydrocarbons have been widely studied in fixed bed reactors, the higher C H 4 / O 2 ratios required to minimize complete oxidation lead to low C H 4 conversions and moderate C 2 yields [4,6,7]. Several reactor configurations and novel approaches have been proposed to achieve higher C 2 yields for the OCM reaction. The approach employed in the Arco cyclic process involved transfer of oxygen-storing solids between oxygen-charging and methane conversion reactors [18]. Recently, a different approach involving product removal, that gave rise to higher C 2 yields ( ~ 50%) for the OCM reaction, was reported by Tonkovich et al. [19] through the use of a staged chromatographic reactor. In this simulated counter-current moving bed chromatographic reactor (SCMCR), methane conversion was significantly increased by introducing intermediate stages of product removal through a moving bed chromatographic column. A similar approach was taken by Jiang et al. [20] who developed a molecular sieve trap in the recycle loop of an electrocatalytic reactor-separator in order to accomplish product removal and improve process yields ( ~ 80%). While these cyclic processes seem attractive because of enhanced C 2 yields, the energy requirements and complexity of the setups pose problems for commercialization. The methane coupling process must operate at high temperatures ( > 750°C), while the adsorption of the desired products in the chromatography columns is favored at low temperatures ( < 200°C). Dautzenberg et al. [21] report that while the ideal coupling reaction itself is exothermic ( ~ 66 kcal/mol of C 2 H 4 ) , the overall contribution of deep

oxidation reactions makes the process even more exothermic ( ~ 123 kcal/mol of C2s produced, based on 84% selectivity to C2s). Their study concludes that multitubular reactors of impractical dimensions would be required to control the heat transfer and avoid temperature runaway, and that newer reactor designs are necessary for commercialization. Clearly, heat transfer issues will continue to pose challenges for membrane reactor development. One of the newer approaches for attaining improved hydrocarbon selectivity in the OCM process relies on the concept of limiting the supply of oxygen necessary for the coupling reaction, which inhibits the deep oxidation (homogenous gas phase reaction pathways) that leads to CO and CO 2 products [1,24]. A lowered oxygen concentration should improve the hydrocarbon selectivity since the apparent reaction order for oxygen in the deep oxidation reaction is higher than in the coupling reaction [22]. Limiting the oxygen supply is attractive since, apart from maintaining conditions conducive to hydrocarbon formation, an additional design parameter is provided for thermal control of the highly exothermic methane coupling reaction. Modeling studies on oxygen metering in shell-and-tube type membrane reactors by Tsai et al. [23] point to the beneficial role of membrane reactors in mitigating the temperature rise in exothermic reactions. Control of feed policy was also studied by Choudhary et al. [24], who proposed fixed bed reactors with staged/distributed oxygen feed. For a fixed overall methane to oxygen ratio, C 2 yield increased as the number of oxygen feed points increased. This approach was also taken by Santamaria and co-workers (Lafraga et al. [25], Coronas et al. [26,27]) through the use of modified alumina membrane reactors, and in the theoretical work of Reyes et al. [22]. These studies compared the performance of the membrane reactor with distributed oxygen feed to a conventional quartz tube reactor with a methane and oxygen co-feed. Lafraga et al. [25] and Coronas et al. [26] showed that the membrane reactors gave higher hydrocarbon selectivities, especially at low and moderate methane conversions. Experimental studies therefore suggest that improvements in C 2 selectivity could be realized in the OCM reactions carried out under metered flow in membrane reactors. The objective of our research was to investigate

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the effect of metering the oxygen supply on the oxidative coupling of methane while carrying out the reaction at high levels of methane conversion. Shelland-tube type membrane reactors were employed in our studies, where the catalyst was packed in porous Vycor membrane tubes. Oxygen flowed uniformly across the length of the catalyst bed, from the shellside to the tube-side through the porous Vycor membrane, at a rate controlled by the applied pressure differential between the shell and the tube. Our studies are unique in that the use of thick-walled porous Vycor membranes enabled a more uniform delivery of oxygen down the length of the catalyst bed.

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2. Materials characterization

Several commercially available porous materials have been investigated for use as membranes to regulate the flow of oxygen in the membrane reactor. Porous Vycor of pore size 40 A was chosen in this work, due to the fact that controlled calcination of these membranes permitted the gas permeance to be adjusted to values appropriate for studying methane coupling in our laboratory setup. A modified porous Vycor membrane can withstand the high temperatures required for the methane coupling reaction and also provide a steady oxygen flux, allowing better distribution of the reactant gases over the catalyst bed than in a conventional plug flow reactor, thereby maintaining a low gas phase concentration at the catalyst site. The 1.5 mm thick walls of the 40 A pore diameter porous Vycor tubes caused very low values of diffusive flux across the membrane. Oxygen flow rates were typically about 10 ml/min across the entire porous Vycor tube when a pressure differential of about 60 psi was applied across the membrane. Fig. 1 shows the variation in helium and oxygen permeate flow through the porous Vycor membrane as a function of the pressure differential applied across the membrane wall. The oxygen flow could thus be easily varied in the range conducive to experimental studies in the laboratory scale setup. Moreover, since the pressure drop due to the catalyst packed inside the Vycor tube was only 1 to 2 psi, the oxygen flux was essentially unchanged along the length of the catalyst bed.

0

100

200

300

Pressure drop across membrane

i 400 (kPa)

' 500

Fig. 1. Variation in helium (+) and oxygen (0) permeate flow across porous Vycormembrane with applied pressure differential across the membrane.

The starting materials for the modified membranes were commercially available porous Vycor tubes. Tubes of dimensions 10 mm o.d., 7 mm i.d. and 100 cm in length were purchased from Coming Inc. (code 7930). The membranes had an average pore diameter of 40 A with a void fraction of about 28%. Gas permeance of nitrogen through these 40 ,~, 1.5 mm thick, porous Vycor tubes was measured to be on the order of ~ 1 0 - 9 m o l / m 2 s Pa, at room temperature. The Vycor tubes were pre-treated at various temperatures before use. The pre-treatment procedure involved calcination of the Vycor tubes at a fixed temperature for 48 h. Gas permeability through the membrane decreased rapidly with increasing pre-treatment temperature between 850 and 1000°C. Though there was a significant change in gas permeability through the membrane in this temperature range, at steady-state at a fixed temperature, this permeability was constant. At a constant operating temperature below the pre-treatment temperature, gas permeability through the membrane was found to remain steady over extended periods of time. Since a typical experimental run would last between 24 and 48 h, and would be carried out at temperatures lower than 850°C, all the Vycor membranes were pretreated at 900°C for 48 h before use in each experi-

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3. Reactor design and experimental details

ment. The fact that gas permeance through the membrane could be adjusted by controlled calcination, which resulted in controlled pore closure, was especially beneficial to our study and helped in devising a reliable membrane module fabrication procedure. Shelekhin et al. [28] reported a thorough study of heat-treated Vycor glass. They found that the permeability of helium gas through porous Vycor underwent a sharp decrease at pre-treatment temperatures greater than 850°C, in a very similar fashion to the present results with nitrogen. They observed a linear shrinkage of about 2.5% at 850°C pretreatment temperature, which was confirmed in the present study. The present work also found identical dependences of pore volume (cm3/g) and BET surface area (m2/g) with the pre-treatment temperature, as were reported by Shelekhin et al., who concluded that this was possible only if there were no change in the pore diameter with pre-treatment temperature. Therefore, the majority of the pores in porous Vycor that did not collapse with the thermal pre-treatment were still of 40 A pore diameter. The variation in permeability was therefore not a consequence of a decrease in pore size, as was the case in the membrane reactor studied by Coronas et al. [26,27]. The constant pore size throughout the porous section of our reactor minimized any potential radial and axial inhomogenities in the amounts of oxygen permeate flow along the porous section of the membrane module.

v

A composite five-section Vycor-quartz membrane module served as the tube of the shell-and-tube reactor. Fig. 2 shows a schematic of the five-section, porous Vycor/non-porous quartz, composite membrane module. The membrane module consisted of a well-defined central porous region with a permeability to nitrogen of 20 000 Barrer, porosity-minimized sections on either side of the central porous region with permeability less than 200 Barrer (essentially non-porous relative to the porous section), and two non-porous quartz tube end sections. On the basis of the results of material characterization, the following fabrication procedure was devised. A 30 cm-long porous Vycor tube was first pre-treated at 900°C for 24 h. The porosity-minimized sections were then made from this pre-treated tube by subjecting ~ 10 cm from the two ends to an additional thermal treatment at 1100°C for 24 h. The Vycor tube was then cut so that the porous section was centered. The two end sections were then glass-welded onto two non-porous quartz tubes of 10 mm o.d. and of equal length to make up a five-sectional membrane module with a total length of 52 cm. The length of the porous section could be varied by subjecting different lengths of the Vycor material to the higher temperatures, but in a typical membrane module, the porous section was about 10 cm and the porosity-re-

02

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~2

Fig. 2. Schematicof longitudinalcross-sectionof membranemodule(no thermowell).

A.M. Ramachandra et al. / Journal of Membrane Science 116 (1996) 253-264

duced sections were about 7 cm. The transition region between the porous central region and the porosity reduced Vycor was generally limited to less than 1 cm. The membrane module thus fabricated was housed inside a quartz-lined stainless steel tube to make up the membrane reactor. The porous membrane reactor was a shell-and-tube type reactor with the catalyst packed inside the membrane tube. The membrane module which served as the tube for the membrane reactor, or the quartz non-porous tube which was used in the conventional co-feed reactor, was fitted into the reactor by using Kalrez ® O-rings to provide leak-tight seals between the tube and the shell. The catalyst, mixed with quartz chips, was confined to the porous region of the membrane tube (or the center of the non-porous quartz tube) and was packed in the annular region between a 3 mm o.d. quartz thermowell and the 7 mm i.d. tube. Fig. 3 shows a cross-sectional view of the annular catalyst bed in the shell-and-tube membrane reactor. The rest of the tube side volume was packed with quartz chips for the purpose of reducing the reactor volume and thus minimizing non-selective gas phase reactions. The catalyst that was selected for the experimental study was samarium oxide (Sm203). Apart from being the most commonly-used methane coupling

10 m m OD Vy( m e m b r a n e tuk

257

catalyst [1,17], and commercially available in high purity grade, samarium oxide also showed no deactivation and loss in selectivity during more than 100 h on-stream. Since the experimental thrust was on comparing the reactor performance of the membrane reactor and the conventional co-feed reactor, the use of a simple, stable oxide obviates the problems of catalyst dispersion (as for L i / M g O [22-26]) or concentration inhomogenities (as for mixed oxide catalysts such as SrCeYbO 3 [11,15]). Sm20 3 (particle size ~ 1-2 /zm) was purchased from Aldrich Co. (cat. no. 22,867-2). The catalyst bed was calcined at 800°C for 8 h in flowing helium prior to the start of the catalytic studies. A quartz thermowell was positioned along the axis of the membrane tube and allowed a sliding thermocouple to profile the temperature along the reactor bed. The presence of the thermowell not only helped to minimize oxygen starvation at the core of the catalyst bed by creating an annular bed, but also allowed the sliding thermocouple to profile the axial temperature while the reactor was operating at high temperature. Fig. 4 shows the temperature profile along the axis of the shell-and-tube membrane reactor. A three-sectional cylindrical heater system with individual PID controllers maintained near-isothermal conditions (1.2% temperature deviation) over the

.........

rtz thermowell

D Annular t bed 15 m m ID Qua=,,=- ............

Fig. 3. Cross-sectionalview of the catalystpackedmembranereactor.

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12 cm-long porous membrane-catalytic bed section of the membrane reactor. Cooling water lines around the reactor body on either side of the heat zone maintained the O-ring fittings of the shell-and-tube membrane reactor at room temperature. Methane gas diluted with five times as much helium was fed to the tube side. Pure oxygen was fed to the shell side and its flow into the tube was controlled by the pressure differential between the shell and the tube sides, set by using back pressure controllers. Feed flow was controlled using mass flow controllers (models 5850 series C/5851 series E) which were purchased from Brooks Instruments. Argon gas flow was regulated through an independent MFC and was mixed with either the tube side or the shell side effluent in a spiral-tube mixer before being sent to the GC for analysis. The reactor piping

Pre-heater

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Main Heater

II

Catalyst Bed

Post-heater

II

I

Tube

/

700

¢0

was 1 / 8 inch SS tubing except at the entrances and exits to the shell-and-tube reactor which were 1 / 4 inch SS. The entire post-reactor piping, including the GC sampling loop, was kept at 150°C using heating tapes and insulation. A blow-out patch and a high pressure relief valve were included in the reactor circuit on both the tube- and the shell-side inlet streams. A schematic of the experimental setup is shown in Fig. 5. The shell-and-tube reactor housing either the porous membrane tube or the non-porous quartz tube was brought up to reaction temperature at a rate of 2°C/min, under flowing helium. Once the operating temperature was reached, the shell side was pressurized with pure oxygen. Oxygen permeate flow into the tube side was monitored by measuring the O 2 content in the tube side effluent through G C / M S

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Distance down the reactor tube (cm) Fig. 4. Axial temperature profile down the reactor tube.

60

259

A.M. Ramachandra et al./ Journal of Membrane Science 116 (1996) 253-264

analysis. Methane flow was then introduced by mixing with helium flowing on the tube side, and both the tube- and shell-side exit flows were analyzed for product concentrations. In the case of the co-feed quartz tube reactor, first the oxygen, and then the methane, were mixed in appropriate amounts with the helium gas and were co-fed into the tube, and the tube-side exit flow was analyzed. The effluent flows from the reactor were analyzed by an on-line GC-MSD system (HP5890 series II gas chromatograph, HP 5971 mass selective detector). A Poraplot-Q megabore column (0.53 mm column ID, purchased from Alltech Co.) was used to resolve the C H 4 , C 2 H 4 , C 2 H 6 , and CO 2 peaks, while the Ar, CO and O 2 peaks that co-eluted were resolved by the mass spectrometer detector. Concentrations of C H 4 , 0 2 , C 2 H 4 , C 2 H 6 , C O and C O 2 w e r e computed by calibrating against a standard gas mixture containing all the product gases in known quantities, and CH 4 conversions and C 2 ( C z H 4 + C z H 6) selectivities and yields and the carbon balance were calculated. In order to compare the performance of the Vycor

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Fig. 5. Schematic of the membrane reactor experimental setup.

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CH41 0 2 Ratio Fig. 6. Effect of changing C H 4 / O 2 ratio on the performance parameters during oxidative coupling of methane in Vycor membrane reactors ( + , CH 4 conversion; D, C 2 selectivity; A, C 2 yield).

membrane reactor with the co-feed reactor, experimental studies of the oxidative coupling of methane employed a methane to oxygen ratio of two. Literature studies indicated that a conventional packed-bed, co-feed reactor provided maximum hydrocarbon yields while operating at a C H 4 / / O 2 ratio of 2.0 to 3.0 [1,5,25,26]. A higher ratio would decrease overall methane conversion, while a ratio lower than 2.0 would reduce C 2 selectivity. In the case of the Vycor membrane reactors, a C H 4 / 0 2 ratio of 2.0 also maximized the C 2 yields, as can be seen from the plots of CH 4 conversion and C 2 selectivity and yield shown in Fig. 6. Since both the membrane reactor and the non-porous reactor exhibit good hydrocarbon yields at a methane to oxygen ratio of 2.0, feed flow rates used in these comparison experiments were based on a C H 4 / O 2 ratio of 2.0 to avoid undue advantage to either reactor type. Blank reactor experiments were carried out, in which the reactor was filled with quartz chips and equivalent flow and temperature conditions were maintained as for the catalyst-packed membrane and co-feed quartz tube reactors. For both the blank reactor runs as well as the catalytic runs, closure was obtained on the carbon balances between 96 and

A.M. Ramachandra et al. / Journal of Membrane Science 116 (1996) 253-264

260

102%. For the lowest reactant flow rates, methane conversions in the blank runs were in the range of 7 - 9 % for the Vycor membrane reactor, and 4 - 5 % for the blank runs in the quartz non-porous reactor, both at 800°C. At higher reactant flow rates, the blank conversions were much lower (2-3%). The reactor performance parameters, CH 4 conversion, C 2 selectivity and C 2 yield, were calculated based on outlet compositions according to the following definitions: CH 4

40

35

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o

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conversion

: {[moles(CO 2 + CO)] + 2[moles(CzH 4 + C2H6)] } × {[moles(CO 2 + CO)]

15

10 600

650

700

750

800

850

Temperature (°C)

+ 2[moles(C2H 4 + C2H6) ] + [moles CH 4 ]} -1

Fig. 7. Comparison of C 2 selectivities from porous Vycor membrane reactor ( • ) and co-feed, non-porous quartz reactor ( + ).

C 2 selectivity = 2[moles(C2H 4 + C2H6) ] × {[moles(CO2 + CO)] + 2[moles(C2H 4 + C2H6)] } - ' C 2 yield = 2[moles(CzH 4 + C2H6) ] x {[moles(CO 2 + CO)] + 2[moles(C2H 4 + C2H6)] + [moles CH 4 ]} -]

4. R e s u l t s and d i s c u s s i o n

Comparisons of reactor performance were made between the porous Vycor membrane reactors and the conventional non-porous packed bed reactors packed with Sm203 catalyst and operating under equivalent experimental conditions of reaction temperature, catalyst loading, and feed flow rates. The catalyst bed length was 12 cm with a catalyst (Sm203) loading of 2.4 g. Feed flow rates were helium 200 ml/min, methane 40 m l / m i n and oxygen 20 ml/min. For the membrane reactor the oxygen flow was permeated under a shell side pressure of 65 psig. Fig. 7 compares the C 2 selectivities in the porous

Vycor membrane reactor and the quartz non-porous reactor operated at the same contact time of 0.86 seconds. C 2 selectivities were consistently higher in the Vycor membrane reactor over the temperature range 700-800°C, compared to the non-porous quartz tube reactor. With increasing temperatures, the nonselective gas phase methane conversion to CO x products increased. Hence the C 2 selectivity decreased as the reaction temperature increased over 750°C. The overall C 2 yields were approximately the same in both reactor configurations, as shown in Fig. 8. Total methane conversions for the two reactor systems are shown in Fig. 9. The higher conversions seen for the quartz tube fixed bed reactor are mainly caused by the deep oxidation to CO x products, as evidenced by the low values of C 2 selectivity at the corresponding temperatures for this reactor configuration. Fig. 10 shows a conversion versus selectivity plot that compares the performance of the Vycor membrane reactor with that of the co-feed non-porous reactor. At the same level of conversion, membrane reactors produce higher C 2 selectivity compared to non-porous co-feed reactors. Since the membrane reactor had inherently lower methane conversions at the higher operating temperatures, for this comparison of conversion vs. selectivity, the data for the quartz tube non-porous reactor corre-

A.M. Ramachandra et al. / Journal of Membrane Science 116 (1996) 253-264

A

¢f

o

I

I

600

t

i

650

700

750

800

850

Temperature (°C)

Fig. 8. Comparison of C 2 yields from porous Vycor membrane reactor ( • ) and co-feed, non-porous quartz reactor ( + ) .

spond to much lower operating temperatures than the membrane reactors. For both the Vycor membrane reactor as well as the quartz non-porous reactor, Fig. 10 shows that higher selectivities were observed at high methane conversions. This was because the activation energy

261

of C 2 formation is much higher than that of CO x formation [1]. At temperatures lower than 700°C, a predominant amount of methane conversion led to CO x formation and C 2 selectivity was negligible. Since, in the membrane reactor, methane combustion was accomplished under limited oxygen supply, the level of methane conversion was not as high as in the co-fed quartz tube reactor. Higher methane conversions require larger contact times in porous membrane reactors. For the same feed flow conditions, longer contact times translate into larger reactor volumes. A larger bed cross-section is not desired as that would increase methane and hydrocarbon losses to the non-selective reactions on the shell side. This problem was addressed by employing a long catalyst bed (10-12 cm) and fast shell-side gas rates. By using a thermowell and quartz chips to dilute the catalyst bed, hot spots and thermal runaways were minimized. The catalyst powder (1-2 /zm) was well dispersed among the 60-80 mesh size quartz chips. The temperature profile shown in Fig. 4 remained unchanged when the feed gases (methane and oxygen) were introduced, indicating that the exothermicity of the reaction was adequately handled and the heat generated during the reaction was efficiently dispersed.

40-

g

_~

3o-

gl c o

20-

o

,it

o

10-

0-

0

20 600

650

700

750

800

850

Temperature (°C)

Fig. 9. Comparison of CH 4 conversions from porous Vycor membrane reactor ( • ) and co-feed, non-porous quartz reactor

(+).

24

28 32 CH4Converslon (%)

36

40

Fig. 10. Conversion vs. selectivity plot for comparing the performance of the porous Vycor membrane reactor ( A ) and the co-feed, non-porous quartz reactor ( + ). The different conversions were obtained at different reaction temperatures.

262

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Use of the 40 A pore diameter, 1.5 mm thickwalled Vycor membrane enabled a more uniform cross-flow of oxygen (flow across the membrane wall) down the catalyst bed as compared to the work of Coronas et al. [26], whose modified alumina membrane reactor setup had a steep axial gradient of oxygen flow. The low value of gas permeance through the Vycor membrane implied that a high pressure differential (60 psi) was needed to obtain a permeate flow of about 10 m l / m i n of oxygen. Since the quartz chip dilution of the catalyst bed created only 1-2 psi upstream head-pressure, a reasonably constant pressure differential was maintained along the bed length, which meant a uniform distribution of oxygen permeate along the bed length. Comparison of the experimental results with the methane coupling results reported by Coronas et al. [26], who used modified porous alumina membranes in a similar shell-and-tube configuration, shows agreement with the trends of higher selectivity and lower conversion in membrane reactors. Our experiments show a higher C 2 selectivity at equivalent methane conversions as compared to the studies of Coronas et al. at equivalent experimental conditions of W/F = 0.01 g m i n / c m 3. This was probably due to better control on metering the oxygen flow axially down the catalyst bed in our system, provided by the uniform pressure driven flow through 40 A pores of the membrane. Their study also concluded that regulating the supply of oxygen along the length of the packed bed provided higher C 2 selectivity in the membrane reactor as compared to a fixed bed reactor, with comparable C 2 yields for both the reactor configurations. Our results confirm this conclusion. Though it is arguable that the uniform oxygen permeate flow along the reactor length could have been detrimental to C 2 selectivity at the lower end of the reaction zone, where there is a higher concentration of Czs, our modeling studies [29] indicated that imposing a gradation in permeance down the catalyst bed was not expected to cause significant improvements in the overall C 2 selectivity. Recent experimental work reported by Coronas et al. [27], with a gradation in oxygen permeation rate along the catalyst bed, indicates that, although improvements were discernible, the overall improvements in C 2 selectivity and C 2 yield were no more than 1-2 percentage points, in agreement with our results.

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Fig. l 1. Comparison of tube side effluent in Vycor membrane reactor with the overall effluent (tube + shell) of the Vycor membrane reactor ( ~ ) and the quartz tube non-porous reactor ( + ) .

The fact that metering the oxygen flow into flowing methane in the catalyst bed enhanced the C 2 selectivity is further supported, if only the products from the tube side effluent are considered in computing the hydrocarbon selectivity. Although the results from this computation cannot be used to claim that membrane reactors outperform the conventional cofeed reactors, it does strengthen the case for oxygen metering, since it is only on the tube side of the membrane reactor that oxygen is metered. Fig. 11 shows the data o n C H 4 conversion and C 2 selectivity by considering the performance of only the tube side of the membrane reactor. On this conversionselectivity plot, the Vycor membrane tube exhibits higher C 2 selectivity at lower methane conversions than the combined shell-and-tube reactor, which was shown in Fig. 10. These data therefore strengthen the argument of C 2 selectivity enhancements due to oxygen metering, since mainly deep oxidation of methane occurs for the portion that escapes to the oxygen-rich shell side. This indicates that modifications in membrane reactors that minimize losses to the shell side, such as dense membrane reactors that selectively permeate only oxygen, hold promise for further advances in the field of OCM technology.

A.M. Ramachandra et al. / Journal of Membrane Science 116 (1996) 253-264

5. Conclusions In the overall p e r f o r m a n c e comparison, porous V y c o r m e m b r a n e reactors exhibited higher C 2 selectivities at the s a m e c o n v e r s i o n levels, w h e n c o m pared to a non-porous quartz tube fixed bed reactor, indicating that m e t e r i n g the o x y g e n supply along the reactor bed was beneficial for the oxidative c o u p l i n g o f methane. Our w o r k thus demonstrates the advantages of the o x y g e n metering concept for O C M . A l t h o u g h our studies do not s h o w marked imp r o v e m e n t s in the total C 2 yields in V y c o r m e m brane reactors, using S m 2 0 3 catalyst, they clearly indicate that i m p r o v e m e n t s w e r e obtained in C 2 selectivity through o x y g e n metering in m e m b r a n e reactors. These results warrant further investigations into i m p r o v e m e n t s in reactor designs that could result in c o m m e r c i a l i z a t i o n o f the O C M technology. O n e such research avenue is dense m e m b r a n e reactors w h e r e the dense m e m b r a n e s w o u l d prevent methane and h y d r o c a r b o n losses to the shell side, while at the s a m e time permit significant flux o f o x y g e n across the m e m b r a n e .

Acknowledgements The authors w o u l d like to a c k n o w l e d g e the financial support for this w o r k p r o v i d e d by the U S Department o f Energy, Pittsburgh E n e r g y T e c h n o l o g y Center, Pittsburgh, PA, under contract no. DEAC21-92PC92113.

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