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Roles of oxygen and carbon dioxide on methane oxidative coupling over CaO and Sm2O3 catalysts
ELSEVIER
Applied Catalysis A: General 115 (1994) 243-256
Roles of oxygen and carbon dioxide on methane oxidative coupling over CaO and Sm203 catalysts Tang Cailin, Zang Jingling, Lin Liwu* Dalian Institute of Chemical Physics, Chinese Academy of Science, PO Box 1 IO, Dalian, 116023, China
Received 21 October 1993, revised 14 Februaq 1994, accepted 14 April 1994
Abstract The effect of partial pressures of oxygen and carbon dioxide on the rate of methane conversion over CaO and Sm,O, catalysts has been investigated by measuring the concentration gradients of products and reactants along the catalyst bed. The results show that the oxidative reaction of methane mainly occurred in the reaction zone of the bed where the methane, oxygen and catalyst co-exist. In this zone, although the oxygen concentration decreases significantly, the rate of methane conversion remains constant over Sm209 catalyst at 873 K. No inhibition of the catalyst activity by carbon dioxide was observed over Sm20, catalyst. By contrast, at the same reaction temperature of 873 K, CaO catalyst can be completely deactivated behind the reaction zone where the partial pressure of carbon dioxide produced was only 0.8 KPa. When the reaction temperature was elevated to 973 K and 1073 K, the rate of conversion of methane increased sharply to a maximum at the entrance of the bed and then slowly dropped with increasing bed height due to carbon dioxide formation. The curve of rate distribution displays a sharp drop at the point where the concentration of oxygen reaches zero. Based on the above results, a kinetic equation for the methane conversion rate has been suggested which is zero order in oxygen and first order in methane, and the isotherm for carbon dioxide adsorption can be described by a Langmuir equation. Keywords:Calcium oxide; Methaneconversion;Samariumoxide
1. Introduction The oxidative coupling of methane (OCM) is a method for the utilization of natural gas as a raw material for the synthesis of ethylene. Recently significant advances in the reaction mechanism have been proposed [ 1,2]. It has been widely accepted that the reaction involves the initial formation of methyl ( CH3 - ) radicals which undergo coupling to form ethane, with further oxidation of ethane producing *Correspondingauthor.Tel. ( + 86-4 1I ) 3645007, fax. ( + 86-4 1I ) 3632426. 0926-860X/94/$07.00 0 1994 Elsevier Science B.V. All rightsreserved SSDI0926-860X(94)00086-7
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ethylene [ 3,4]. Due to the observation of an isotropic effect (Ku/K,,) over many catalyst systems, the initial step in the methane oxidation to produce the methyl radical has been identified to be the rate-determining step in the reaction [ $61. Although the evidence for the involvement of methyl radicals in the formation of the C2 products is conclusive [ l-61, the details of the mechanism, such as the role of gaseous oxygen, the inhibition of the reaction by carbon dioxide are still unsolWA.
In recent years the effect of carbon dioxide on methane conversion has been extensively studied [ 7-l 5 ] . The inhibition of methane conversion by carbon dioxide has been found in many catalyst systems, especially in the alkaline earth oxides MgO [ 111, CaO and SrO [ 141. Although it is reasonable to attribute the inhibition effect to the formation of surface [CO:- 1, the extent of the inhibition and its mechanism are not yet very clear. The inhibiting effect of carbon dioxide depends on the concentration of the carbon dioxide produced. Since the concentration of oxygen varies along with the height of the catalyst bed, attention has to be focused on the interior of the catalyst bed rather than only the catalyst body. However most results concerning kinetic analysis consider the catalyst bed as a body and the reaction behavior in the interior of the catalyst bed has been ignored. A flowing bed reactor has normally been used to characterize the OCM reaction. Under these conditions, the concentration of gaseous species and catalyst activity at different bed heights have to be kept constant. If this concentration distribution can be measured along the whole catalyst bed, the methane conversion and selectivity distribution can easily be obtained and further information about the reaction mechanism can be derived. In this study concentration distributions along the catalyst bed have been measured in situ by means of a high sensitive mass spectrometer, with a special designed sampling system. Results are presented for the application of this method to the samples of CaO and Sm*O,. Measurements of the concentration distributions were performed to determine the methane conversions, the amount of catalyst surface active sites and the product selectivity along the catalyst bed. The qualitative relationship of these factors with the carbon dioxide and oxygen concentrations in the gas phase have also been established.
2. Experimental 2.1. Principle The general principle underlying the measurement of the concentration distribution is based on the fact that when the OCM reaction is run at steady state in a flowing reactor system, the concentration of every gas component as well as the activity of the catalyst will remain constant at different bed heights. In principle, if the concentration distributions can be measured, both reaction and catalyst per-
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formance in the reactor can be derived according to the following formulas ( 1) and (2) [16]: X(h) =
G,,(h) C,,,(h)
+G,,,(h)
+ 2C,,,,(h)
R(h) = VdX(h)ldh
+G,,,(h)
+ C,,(h)
+ C,,(h)
+ C,,(h)
+ C,,(h)
(1) (2)
where, Cc,,,(h) , CCzHa(h) , Cc,(h) , Ccoz( h), Cc,(h) denote the concentration of C,&, &I-&, CO, CO*, and CH4, respectively, X is the methane conversion and R is the reaction rate. All of the parameters are a function of bed height; V is the flowing velocity of the feed. It is maintained constant within a reactor bed. The concentration distribution can be measured by analyzing the gas samples collected in situ from the catalyst bed at different heights. The question is how to collect the gas sample at high temperature ( > 873 K) without disturbing the stability of the concentration distribution. 2.1. The sampling system A quartz sampling tube with an internal diameter of 1 mm and outer diameter of 2 mm is located at different heights of the bed, one end of this tube is a micro-leak for sampling, another end is directly connected to the vacuum chamber of the mass spectrometer (see Fig. 1) . While the reaction is running, the sampling tube can be moved axially in the bed. Through the micro-leak a limited amount of the gas sample around the leak is continuously admitted to the vacuum chamber and analyzed by the MS detector. The size of the leak has to be designed carefully because: First, the leak must have a sufficient resistance to maintain 1 atm ( = 101.3 kPa) pressure difference between the outer and inner of the sampling tube. Second, the size of the leak must be much smaller than that of the catalyst particles to prevent any blocking by fine particles. Third, the amount of sampling must be very small compared to the flow-rate of the feed in order to avoid bothering the reaction stability. However sufficient flow of sampling gas was necessary for rapid detection. The lower the sampling gas flow, the longer the time required for the sample to diffuse from the leak to the vacuum chamber. The sampling flow could be determined by the pressure of the mass spectrometer chamber. To satisfy the above requirements, the size of the leak was adjusted to a chamber pressure of 5 - 10m6 Torr ( 1 Torr= 133.3 Pa) under reactions conditions and the sampling flow was less than 1 ml/min which was estimated by the flow-rate difference of the effluent gas with and without sampling. 2.3. Product analysis A highly sensitive quadrupole mass spectrometer (ANEVAL36OS) was used to analyzed the gas sample. Mass spectrometry (MS) spectra of standard gas samples
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Catalysis A: General 115 (1994) 243-256
Tomassspeurometer
I -
-
Bed height (mm)
.I..
‘x.
x.
‘.... .._ ‘...
Leak for gas i&A.,,
20
o*
co,
10
‘;
Catalyst ___....... 0
a-&O,
............. l&L 0
._/.,_... /..~~’ _ Heating
5
Concentration
(%)
reaction gas Fig. 1. Illustration of the reactor and the sampling system.
regularly scanned from l-70 amu at intervals of 1 mm were used to identify the gas species and their concentrations. The analysis of the spectra at different bed height showed that HZ, C&, 02, CO2 and Hz0 can be identified appropriately by the mass numbers M/Z= 2, 16, 32,44 and 18 respectively. The ion signals of Ml Z= 26 and 30 varied in a way different from that of M/Z= 28, and the ion signal of M/Z= 27 showed the sum of M/Z= 26 and M/Z= 28. So the peak at M/Z= 26 was assigned to C,H,. Similarly, the peak at M/Z= 30 was assigned to C,&. The correct amount of CO was calculated from M/Z= 28 by subtracting the contribution due to C2H,. We ignored the contribution of C& to M/Z= 26 because it was very low compared to the amount of CO. Theoretically, under the conditions of the constant flow resistance of the leak and constant pressure difference between the leak, the amount of a gas sampling into the MS will be proportional to its concentration per volume in the gas near the leak, thus it is also proportional to the intensity of the corresponding MS peak. Eq. (3) was used for quantification of the concentration of gas component. Ci =f( Zi - Zoi)
(3)
where, Ci refers to the concentration of gas component i; Zi corresponds to the ion intensity detected by MS; Zoirefers to the intensity from background; f is a factor determined by a normal gas chromatographic analysis of a standard gas.
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2.4. Reaction conditions
The CaO and Smz03 catalysts were prepared by crushing the compounds (their purity 99.5%) to 30-60 mesh, followed by calcining in air for 5 h at 1073 K. The catalytic reaction was performed at atmospheric pressure using a conventional quartz reactor with an internal diameter of 6 mm, as shown in Fig. 1. Such a reactor, when filled with cr-Al*O, and exposed to the reaction mixture at a temperature of 1073 K, gave only moderate methane conversion ( < 2%). In a typical experiment, 700 mg of catalyst was loaded into the reactor, supported by inert o-A&O,, and then pre-treated in flowing helium for 2 h at 1073 K. When no carbon dioxide could be measured in the effluent gas by MS, the reactor was cooled to the reaction temperature and the feed was then switched to the reaction mixture of 46% CH,+38%He+16%02forSmzO~and38%CI-L,+11%O02+51%HeforCaO. The flow-rate of the reaction mixture was 46 ml/min. In all experiments, high purity U-I.+(99.5%), O2 (99.5%) and He (99.995%) were used. A complete test ( > 8 h) of the catalyst was performed in the conventional temperature range of 600-1073 K over CaO or SmZOs and the results showed that the OCM performance of these catalysts was very stable under these reaction conditions. After an initial time of 2 h, methane conversion and CZ selectivity did not change. These results ensure that the measurement of concentration bed distribution is acceptable, when performed at different times.
3. Results and discussion 3. I. Methane conversion along catalyst bed Methane conversion at 873 K over Smz03 catalyst as a function of the bed height is given in Fig. 2. In this figure, a height of zero represents the interface of the catalyst and supporting ar-Alz03. The concentration distributions of oxygen and carbon dioxide are also shown in Fig. 2. Temperature measurements over the working catalyst did not show a significant gradient ( < 5 K).Over Sm203 catalyst the reaction mainly occurred in the region ranging from zero to 11 mm from the front of the bed. Within this region, where the methane, oxygen and catalyst coexist, the methane was oxidized by oxygen species of the catalyst. Behind this reaction zone, no further conversion of methane was observed. In order to distinguish this part of the bed from the entire catalyst bed, we specified this region as the reaction zone of the bed. Within the reaction zone, the methane conversion increased linearly with bed height to a maximum at the end of the reaction zone and then remained constant. The linear increase of methane conversion as a function the bed height reveals that methane was oxidized with a constant reaction rate within this region. If the concentration of the active site on Sm,O, catalyst surface
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Catalysis A: General 115 (1994) 243-256
25 16
6
0
0 5
0
10
15
20
Bed Hciglu (nm) Fig. 2. The methane conversion and the concentration distributions of O2 and CO, along the Sm,Og catalyst bed at 873 K.
is assumed to be [ Oact] , methane is oxidized over these sites to form methyl radical and surface hydroxyl: CT& + [ Oact] -
Kl
CH3 * + [HO]
(I)
The reaction rate of the above reaction in fixed bed reactor can be expressed as
[I61 R=
VdX
-
dh
=
& pat400
(4)
where dXldh represents the reaction rate and corresponds to the slope of the methane conversion curves as a function of the bed height, PC+, is the concentration of methane in the flowing gas, r3o is the effective concentration of surface active oxygen [ Oact] when the catalyst is exposed to methane. The parameter K, 0, referred to the activity of the catalyst when it is working in the reactor. The variation in this parameter along the reaction zone can be calculated from the concentration distribution. As shown in Fig. 3, the activities of the working catalyst along the whole bed remain essentially constant. A comparison of the catalyst activity or the reaction rate with the concentration of oxygen and carbon dioxide along the catalyst bed (Fig. 2 and 3) indicates that 6Jodoes not vary with
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the change of gas concentration of oxygen and carbon dioxide. This result is consistent with those found previously for the Sm/CaO system [ 121, where the author found that the rate of methane conversion was also constant, by raising the reaction temperature to 1073 K to desorb carbon dioxide from the Sm/CaO surface. It is also expected on the basis of Langmuir adsorption, that the chemisorption of oxygen over Sm20, catalyst is saturated at low partial pressures ( <0.6 Pa) of oxygen. Results of Amorebieta and Colussi [ 17,181 gave support to this assumption, since the oxygen chemisorption reached saturation with less than 100 mPa of oxygen over Li/MgO [ 171 or Sm*O, [ 181 catalysts. When the reaction was carried out at 973 K over CaO catalyst, the conversion rate of methane and the oxygen concentrations are shown in Fig. 4. At the interface of the a-A1203 and the catalyst, the rate of methane conversion increases sharply, suggesting that the catalyst plays an important role in the conversion of methane. In the reaction zone (O-l 1.5 mm), the oxygen concentration decreases rapidly from the feed concentration of 11% to zero. The reaction rate dropped suddenly at a bed height of 11.5 mm, just at the point where oxygen was completely consumed. It is significant to note that as the gaseous oxygen reaches zero concentration at the end of the reaction zone, the rate of methane oxidation is still maintained at a high level (0.01 mm- ‘) . The same result has also been observed at 1073 K, while a higher reaction rate for methane conversion of 0.05 mm- ’ was observed. This sharp decrease in reaction rate indicates that neither the steady-state adsorption of oxygen nor a Langmuir-Hinshelwood mechanism involving unsaturated adsorption
l.O I t 0.8
.
0’
.
.
.
.
.
l
4k
3
‘3 0.6 e 1 = a v, 0.4 A
0.2
0.0
-
-
-
2
0
2
4
6 Bed
8
10
12
Heigd (mm)
Fig. 3. The catalyst activity ICC&,distribution along the Sm,O, catalyst bed at 873 K.
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10
5
0 -6 -4 -2
0
2
4
6
8
10 12 14 16 18 20 22
Bed Height(xm) Fig. 4. The methane conversion at 973 K.
rate distribution
and the. O2 concentration
distribution along the CaO catalyst bed
of oxygen lead to rate expressions compatible with the observed kinetic behavior. From a consideration of the constant activity of Smz03 in the reaction zone and assuming a Langmuir-Hinshelwood mechanism for oxygen adsorption, it may be concluded that oxygen adsorption is saturated under our reaction conditions, either in dissociative form or in molecular form. The concentration of oxygen in the reaction zone is high enough to result saturation of the adsorption at the operating temperature. This results implies that the concentration of active sites on the catalyst surface is not influenced by the variation in oxygen concentration. The decrease in methane conversion rate in the reaction zone over CaO catalyst may be a consequence of a different mechanism. This topic will be discussed below. 3.2. The role of carbon dioxide When the reaction was performed over CaO catalyst at the same temperature of 873 K as that used for the SmzO, catalyst, the CaO was found to be less active than Sm203. The rate of methane conversion and carbon dioxide distribution along the catalyst bed are shown in Fig. 5. The total conversion of methane was about 1.2%. The main products were the total oxidation products carbon dioxide (selectivity <94%) and carbon monoxide. The catalyst essentially behaved as an oxidation
C. Tang et al. /Applied
0.30
Catalysis A: General 115 (1994) 243-256
251 1.0
-
0.25 A -7
b
--_)
0.8
.
% “a 0.20 = $ d e 0.15 .B E 8 0
;
0.10
1 3
0.2
0.05
0.0
0.00 -2
0
2
4
6
8
10
12
14
16
Bed Height @nn) Fig. 5. The methane conversion rate distributionand the CO2concentrationdistributionalong the CaO catalyst bed at 873 K.
catalyst. Trace amounts of hydrocarbon products (C,H, and &I&) were also detected. Their concentrations along the catalyst bed varied in a similar way to that of carbon monoxide, but were estimated to be less than one tenth of that of carbon monoxide. At the beginning of the reaction zone, the rate of methane conversion increased rapidly to a maximum and then deciined along with the bed height. At a bed height of 9 mm, the reaction rate was essentially zero, while 11% oxygen and 38.6% methane were still present in the gas phase. This results indicates that the catalyst particles located behind the reaction zone are strongly deactivated by carbon dioxide, since large amounts of the reactants, methane and oxygen, co-existed with the catalyst. It is obvious that most of the active sites are occupied by some species which inhibit the adsorption and reaction of methane. In order to find which species play such an important role, temperature-programmed desorption (TPD) of carbon dioxide was performed after the reaction. Before TPD was performed, the reactor was flushed with helium to remove species that could easily desorb. The reactor was then cooled to ambient temperature. It can be seen from the TPD profile of Fig. 6 that a large amount of carbon dioxide was detected, desorption of carbon dioxide commenced at 873 K and reached a maximum at 973 K (shown in Fig. 6). Fig. 6 indicates that a large quantity of carbon dioxide was adsorbed on the catalyst surface. Adsorption of carbon dioxide blocks the surface sites and inhibits the methane conversion. This provides a clear
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4
600
1000
800 T-i=-=
1100
1300
(K)
Fig. 6. TPD spectrum of CO1 over CaO catalyst after OCM reaction at 873 K.
explanation for the decrease in the methane conversion rate of CaO. Initially there was no carbon dioxide in the feed gas, so the surface active sites are all exposed to methane and the reaction rate reached its maximum. As carbon dioxide was formed by the reaction, it covered part of the actives sites and reduced the activity of the catalyst. At a certain height of the bed, the concentration of carbon dioxide was high enough to cover all the surface active sites, and prevent the methane molecule to contact the active oxygen sites, thus resulting in the complete deactivation of the catalyst for methane oxidation. In Fig. 3, it is shown that the rate of conversion of methane remained constant over the whole bed of Sm20, catalyst even when carbon dioxide was present. TPD of carbon dioxide over this catalyst does not show any desorption. This results partly agrees with the results of Peil et al. [ 131 and Dingjun et al. [ 11I. They found that carbon dioxide has less inhibition effect on methane conversion over Sm203 than that over Li/MgO catalyst. Dingjun et al. [ 1l] recently reported that CH3 radical formation over Sm203 is affected much less by carbon dioxide than over Li/MgO catalyst. The carbon dioxide inhibition over CaO and other basic catalysts as MgO, SrO, CaO has also been observed by many groups [ 14,151, however, none observed the significant effect of produced carbon dioxide on active oxygen sites. 3.3. The rate equation The results of concentration distribution measurements provide a basis for quantitatively relating the activity of catalyst with the real concentration of carbon
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Catalysis A: General 115 (1994) 243-256
253
dioxide, according to the reaction scheme. As mentioned above, the active sites were partly covered by produced carbon dioxide, COZ + Oact + [ COJ]
(II)
Under our special experimental conditions, the carbon dioxide adsorption is in equilibrium with the gas-phase concentration at a specific height of the catalyst bed. If the adsorption can be described by a Langmuir relationship, the coverage of COZ ( 0,,) on active sites follows:
632 =
K’PCO2
(3
1 + K’Pco2
where, Pcoz is the partial pressure of CO2 at bed height h, K’ is the equilibrium coefficient. In the previous discussion, we found that the oxygen concentration in the gas phase does not affect the reaction rate. As a result, the concentration of surface active sites in the reaction bed without carbon dioxide coverage can be assigned a value of unity. Therefore the effective concentration of active sites can be expressed as
00=1-ec,=
1 1 + K’Pco2
Finally, the methane conversion rate can be expressed as: R=
WC
-
dh
=
KIPcH., 1+ K’P,,
or
In the reaction, K1, K’, and V were constants, PcHI, PC&, and dX/dh can be obtained from the concentration distribution. It should be noticed that PC&, Pco, referred to the actual concentration over the working catalyst surface, which differs from that obtained from GC analysis in the effluent gas from the reactor. Since conversions at 873 K were very low, results at 973 K and 1073 K only are exhibited in Fig. 7. The results showed that PcH4/ ( dX/dh) varied linearly with carbon dioxide gas concentration. This correlation confirms the validity of Eq. (7) for quantitative expression of the inhibition of carbon dioxide. The adsorption equilibrium coefficient K’ can be obtained from the slopes of the lines in Fig. 7. The values of K’ at 973 K and 1073 K are equal to 0.487 Kpa-’ and 0.306 KPa-‘, respectively. At 973 K, the carbon dioxide concentration at the reaction zone is 5.15 KPa, the effective concentration of surface active oxygen 0, is 0.29. This value implies that 71% of the active sites are covered by produced carbon dioxide. At 1073 K, only 40% of the active sites were exposed to methane at the end of reaction zone. The
C. Tang et al. /Applied Catalysis A: General I15 (1994) 243-256
254
1
0
I,,, 0
(
,,,,,,,,,,,(,,i
1
!
,/,,,,,,,.
2
3
‘I
5
6
Pm2 SW
Fig. 7. PC,, / (dXldh) versus CO2pressureP c- in the reactionzone of the CaO catalyst.
result indicate that the produced carbon dioxide plays a very important role in methane oxidation over CaO catalyst or a similar basic oxide. To study the OCM reaction over this type of catalyst, the role of carbon dioxide must be carefully considered. It has been pointed out that carbon dioxide plays a role as a poison for the oxidative coupling reaction and increases the activation energy for the conversion of methane [ g-l 11. Apart from Roos et al. [ 71 who have considered carbon dioxide as a competitive adsorbate with methane, the inhibiting effect of carbon dioxide has not been considered in kinetic equations for OCM [ 19-231. Based on the results of normal catalytic fixed-bed reactors, many researchers showed that the reaction rate was related to the oxygen concentration in the feed stock, which seems in conflict with our results. It is worth to note that these results considered the catalyst bed as a whole. In our experiments, however, the oxygen concentration and the methane conversion are related directly to different layers of the working catalyst bed. These results have been obtained in one reaction run rather than in many experiments. This method will cause much less temperature variance on the working catalytic surface. In order to avoid severe temperature gradients within the catalyst bed and maintain the oxygen conversion rate below 20%, limited amounts of catalyst loading with large volumes of diluted inert mate-
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255
rial are normally employed at high temperature ( > 973 K).Since a large amount of reaction heat was produced in a small amount of catalyst with a low surface area, the surface temperature of the catalyst is expected to vary extensively in different oxygen conversions. In order to prevent local heating of the catalyst, a large amount of diluent gas in the feed has usually been used. During these experiments, the oxygen concentration is rather low in comparison to our reaction conditions, thus, an unsaturated adsorption of oxygen is expected. As the reaction is run at low temperature, inhibition of the reaction rate by carbon dioxide would be very serious. In some cases, such as CaO, the catalyst behind the reaction zone exhibits no activity for methane oxidation. As a result, the conversion rate per gram of catalyst must be used carefully to express the reaction rate.
4. Conclusion The concentration profile of products and reactants as a function of distance along the catalyst bed has been studied for the OCM reaction over CaO and Sm203 catalysts by means of a specially designed system. It is found that the reaction mainly occurred in the reaction zone of the bed where the oxygen still exists in the gas phase. The concentration of active sites of CaO catalyst is not affected by the concentration of oxygen in the gas phase of the catalyst bed. The active sites of CaO catalyst can be deactivated by the carbon dioxide produced, and this is the main reason for the decrease in the methane conversion in the reaction zone. At low temperature (873 K), the catalyst can be completely deactivated when the partial pressure of produced carbon dioxide was greater than 0.8 kPa. Coverage of carbon dioxide to active oxygen sites can be quantitatively expressed by the Langmuir relationship. In contrast to the CaO catalyst, Sm203 catalyst does not show any inhibition effect by carbon dioxide under the same reaction conditions. All the results showed that the method for measuring the concentration distribution of reactant and product along catalyst bed was successful and effective in the study of the performance of OCM catalysts.
References [ 11 J.H. Lunsford, Catal. Today, 6 (1989) 235. [2] Y. Amenomiya, V.I. Birss, M. Goledzinowki, J. Galuszka and A.R. Sanger, Catal. Rev.-S& Eng., 32 (3) (1990) 163. [3] D.J. Driscoll, W. Martir, J.-X. Wang and J.H. Lunsford, J. Am. Chem. Sot., 107 (1985) 58. [4] P.F. Nelson, C.A. Luckey and N.W. Cam, J. Phys. Chem., 92 (1988) 6176. [S] P.F. Nelson, C.A. Luckey and N.W. Cant, J. Catal., 120 (1989) 216. 161 G. Mirodatos, A. Holmen, R. Mariscal and G.A. Martin, Catal. Today, 6 ( 1990) 601. [7] J.A. Roos, S.J. Korf, R.H.J. Veehof, J.G. Van Ommen and J.R.H. Ross, Appl. Catal., 52 (1989) 131. [ 81 P.F. Nelson and N.W. Cant, J. Phys. Chem., 94 ( 1990) 3756. [9] M. Xu, C. Shi, X. Yang, M.P. Rosynek and J.H. Lunsford, J. Phys. Cbem., 96 (1992) 6395.
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[ 101 S.J. Korf, J.A. Ross, N.A. de Bruijn, J.G. Van Ommen and J.R.H. Ross, Catal. Today, 2 (1988) 535. ] 111 W. Dingjun, X. Mingting, S. Chunlei and J.H. Lunsford, Catal. Lett., 18 ( 1993) 323. [ 121 C. Tang, in L. Guczi, F. Solymosi and P. Tetenyi (Editors), New Frontiers in Catalysis, Proc. 10th Int. Congr. Catal., Budapest, 19-24 July 1992, (Studies in Surface Science and Catalysis, Vol. 75). Elsevier, Amsterdam, 1993, p. 2221. [ 131 K.P. Peil, J.G. Goodwin and G. Mar&in, in A. Holmen, K.-J. Jens and S. Kolbee (Editors), Natural Gas Conversion (Studies in Surface Science and Catalysis, Vol. 61) , Elsevier, Amsterdam, 1991, pp. 73-79. [ 141 T. Suzuki, K. Wada and Y. Watanabe, Appl. Catal., 59 (1990) 213. 1151 X.D. PengandP.C. Stair, J.&ml., 128 (1991) 264. [ 161 C. Tang, Ph.D. thesis, Dalian Institute of Chemical Physics, Academy of Science, People’s Republic of china, 1992. [ 171 B.Z. Amorebieta and A.J. Colussi, J. Phys. Chem., 92 (1988) 4576. [ 181 B.Z. Amorebieta and A.J. Colussi, J. Phys. Chem., 93 (1989) 5155. [ 191 K. Gtsuka and K. Jinno, Inorg. Chim. Acta, 121 ( 1986) 237. [20] E.E. Miro, J.M. Santamaria and E.E. Wolf, J. Catal., 124 ( 1990) 465. [21] E. IwamatsuandK.-I.Aika, J.Catal., 117 (1989) 416. [22] E.E. Miro. J.M. Santamariaand E.E. Wolf, J. Catal., 124 (1990) 465. [23] L. Lehmann and M. Baems, J. Catal., 135 (1993) 467.
Applied Catalysis A: General 115 (1994) 243-256
Roles of oxygen and carbon dioxide on methane oxidative coupling over CaO and Sm203 catalysts Tang Cailin, Zang Jingling, Lin Liwu* Dalian Institute of Chemical Physics, Chinese Academy of Science, PO Box 1 IO, Dalian, 116023, China
Received 21 October 1993, revised 14 Februaq 1994, accepted 14 April 1994
Abstract The effect of partial pressures of oxygen and carbon dioxide on the rate of methane conversion over CaO and Sm,O, catalysts has been investigated by measuring the concentration gradients of products and reactants along the catalyst bed. The results show that the oxidative reaction of methane mainly occurred in the reaction zone of the bed where the methane, oxygen and catalyst co-exist. In this zone, although the oxygen concentration decreases significantly, the rate of methane conversion remains constant over Sm209 catalyst at 873 K. No inhibition of the catalyst activity by carbon dioxide was observed over Sm20, catalyst. By contrast, at the same reaction temperature of 873 K, CaO catalyst can be completely deactivated behind the reaction zone where the partial pressure of carbon dioxide produced was only 0.8 KPa. When the reaction temperature was elevated to 973 K and 1073 K, the rate of conversion of methane increased sharply to a maximum at the entrance of the bed and then slowly dropped with increasing bed height due to carbon dioxide formation. The curve of rate distribution displays a sharp drop at the point where the concentration of oxygen reaches zero. Based on the above results, a kinetic equation for the methane conversion rate has been suggested which is zero order in oxygen and first order in methane, and the isotherm for carbon dioxide adsorption can be described by a Langmuir equation. Keywords:Calcium oxide; Methaneconversion;Samariumoxide
1. Introduction The oxidative coupling of methane (OCM) is a method for the utilization of natural gas as a raw material for the synthesis of ethylene. Recently significant advances in the reaction mechanism have been proposed [ 1,2]. It has been widely accepted that the reaction involves the initial formation of methyl ( CH3 - ) radicals which undergo coupling to form ethane, with further oxidation of ethane producing *Correspondingauthor.Tel. ( + 86-4 1I ) 3645007, fax. ( + 86-4 1I ) 3632426. 0926-860X/94/$07.00 0 1994 Elsevier Science B.V. All rightsreserved SSDI0926-860X(94)00086-7
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ethylene [ 3,4]. Due to the observation of an isotropic effect (Ku/K,,) over many catalyst systems, the initial step in the methane oxidation to produce the methyl radical has been identified to be the rate-determining step in the reaction [ $61. Although the evidence for the involvement of methyl radicals in the formation of the C2 products is conclusive [ l-61, the details of the mechanism, such as the role of gaseous oxygen, the inhibition of the reaction by carbon dioxide are still unsolWA.
In recent years the effect of carbon dioxide on methane conversion has been extensively studied [ 7-l 5 ] . The inhibition of methane conversion by carbon dioxide has been found in many catalyst systems, especially in the alkaline earth oxides MgO [ 111, CaO and SrO [ 141. Although it is reasonable to attribute the inhibition effect to the formation of surface [CO:- 1, the extent of the inhibition and its mechanism are not yet very clear. The inhibiting effect of carbon dioxide depends on the concentration of the carbon dioxide produced. Since the concentration of oxygen varies along with the height of the catalyst bed, attention has to be focused on the interior of the catalyst bed rather than only the catalyst body. However most results concerning kinetic analysis consider the catalyst bed as a body and the reaction behavior in the interior of the catalyst bed has been ignored. A flowing bed reactor has normally been used to characterize the OCM reaction. Under these conditions, the concentration of gaseous species and catalyst activity at different bed heights have to be kept constant. If this concentration distribution can be measured along the whole catalyst bed, the methane conversion and selectivity distribution can easily be obtained and further information about the reaction mechanism can be derived. In this study concentration distributions along the catalyst bed have been measured in situ by means of a high sensitive mass spectrometer, with a special designed sampling system. Results are presented for the application of this method to the samples of CaO and Sm*O,. Measurements of the concentration distributions were performed to determine the methane conversions, the amount of catalyst surface active sites and the product selectivity along the catalyst bed. The qualitative relationship of these factors with the carbon dioxide and oxygen concentrations in the gas phase have also been established.
2. Experimental 2.1. Principle The general principle underlying the measurement of the concentration distribution is based on the fact that when the OCM reaction is run at steady state in a flowing reactor system, the concentration of every gas component as well as the activity of the catalyst will remain constant at different bed heights. In principle, if the concentration distributions can be measured, both reaction and catalyst per-
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formance in the reactor can be derived according to the following formulas ( 1) and (2) [16]: X(h) =
G,,(h) C,,,(h)
+G,,,(h)
+ 2C,,,,(h)
R(h) = VdX(h)ldh
+G,,,(h)
+ C,,(h)
+ C,,(h)
+ C,,(h)
+ C,,(h)
(1) (2)
where, Cc,,,(h) , CCzHa(h) , Cc,(h) , Ccoz( h), Cc,(h) denote the concentration of C,&, &I-&, CO, CO*, and CH4, respectively, X is the methane conversion and R is the reaction rate. All of the parameters are a function of bed height; V is the flowing velocity of the feed. It is maintained constant within a reactor bed. The concentration distribution can be measured by analyzing the gas samples collected in situ from the catalyst bed at different heights. The question is how to collect the gas sample at high temperature ( > 873 K) without disturbing the stability of the concentration distribution. 2.1. The sampling system A quartz sampling tube with an internal diameter of 1 mm and outer diameter of 2 mm is located at different heights of the bed, one end of this tube is a micro-leak for sampling, another end is directly connected to the vacuum chamber of the mass spectrometer (see Fig. 1) . While the reaction is running, the sampling tube can be moved axially in the bed. Through the micro-leak a limited amount of the gas sample around the leak is continuously admitted to the vacuum chamber and analyzed by the MS detector. The size of the leak has to be designed carefully because: First, the leak must have a sufficient resistance to maintain 1 atm ( = 101.3 kPa) pressure difference between the outer and inner of the sampling tube. Second, the size of the leak must be much smaller than that of the catalyst particles to prevent any blocking by fine particles. Third, the amount of sampling must be very small compared to the flow-rate of the feed in order to avoid bothering the reaction stability. However sufficient flow of sampling gas was necessary for rapid detection. The lower the sampling gas flow, the longer the time required for the sample to diffuse from the leak to the vacuum chamber. The sampling flow could be determined by the pressure of the mass spectrometer chamber. To satisfy the above requirements, the size of the leak was adjusted to a chamber pressure of 5 - 10m6 Torr ( 1 Torr= 133.3 Pa) under reactions conditions and the sampling flow was less than 1 ml/min which was estimated by the flow-rate difference of the effluent gas with and without sampling. 2.3. Product analysis A highly sensitive quadrupole mass spectrometer (ANEVAL36OS) was used to analyzed the gas sample. Mass spectrometry (MS) spectra of standard gas samples
246
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Catalysis A: General 115 (1994) 243-256
Tomassspeurometer
I -
-
Bed height (mm)
.I..
‘x.
x.
‘.... .._ ‘...
Leak for gas i&A.,,
20
o*
co,
10
‘;
Catalyst ___....... 0
a-&O,
............. l&L 0
._/.,_... /..~~’ _ Heating
5
Concentration
(%)
reaction gas Fig. 1. Illustration of the reactor and the sampling system.
regularly scanned from l-70 amu at intervals of 1 mm were used to identify the gas species and their concentrations. The analysis of the spectra at different bed height showed that HZ, C&, 02, CO2 and Hz0 can be identified appropriately by the mass numbers M/Z= 2, 16, 32,44 and 18 respectively. The ion signals of Ml Z= 26 and 30 varied in a way different from that of M/Z= 28, and the ion signal of M/Z= 27 showed the sum of M/Z= 26 and M/Z= 28. So the peak at M/Z= 26 was assigned to C,H,. Similarly, the peak at M/Z= 30 was assigned to C,&. The correct amount of CO was calculated from M/Z= 28 by subtracting the contribution due to C2H,. We ignored the contribution of C& to M/Z= 26 because it was very low compared to the amount of CO. Theoretically, under the conditions of the constant flow resistance of the leak and constant pressure difference between the leak, the amount of a gas sampling into the MS will be proportional to its concentration per volume in the gas near the leak, thus it is also proportional to the intensity of the corresponding MS peak. Eq. (3) was used for quantification of the concentration of gas component. Ci =f( Zi - Zoi)
(3)
where, Ci refers to the concentration of gas component i; Zi corresponds to the ion intensity detected by MS; Zoirefers to the intensity from background; f is a factor determined by a normal gas chromatographic analysis of a standard gas.
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2.4. Reaction conditions
The CaO and Smz03 catalysts were prepared by crushing the compounds (their purity 99.5%) to 30-60 mesh, followed by calcining in air for 5 h at 1073 K. The catalytic reaction was performed at atmospheric pressure using a conventional quartz reactor with an internal diameter of 6 mm, as shown in Fig. 1. Such a reactor, when filled with cr-Al*O, and exposed to the reaction mixture at a temperature of 1073 K, gave only moderate methane conversion ( < 2%). In a typical experiment, 700 mg of catalyst was loaded into the reactor, supported by inert o-A&O,, and then pre-treated in flowing helium for 2 h at 1073 K. When no carbon dioxide could be measured in the effluent gas by MS, the reactor was cooled to the reaction temperature and the feed was then switched to the reaction mixture of 46% CH,+38%He+16%02forSmzO~and38%CI-L,+11%O02+51%HeforCaO. The flow-rate of the reaction mixture was 46 ml/min. In all experiments, high purity U-I.+(99.5%), O2 (99.5%) and He (99.995%) were used. A complete test ( > 8 h) of the catalyst was performed in the conventional temperature range of 600-1073 K over CaO or SmZOs and the results showed that the OCM performance of these catalysts was very stable under these reaction conditions. After an initial time of 2 h, methane conversion and CZ selectivity did not change. These results ensure that the measurement of concentration bed distribution is acceptable, when performed at different times.
3. Results and discussion 3. I. Methane conversion along catalyst bed Methane conversion at 873 K over Smz03 catalyst as a function of the bed height is given in Fig. 2. In this figure, a height of zero represents the interface of the catalyst and supporting ar-Alz03. The concentration distributions of oxygen and carbon dioxide are also shown in Fig. 2. Temperature measurements over the working catalyst did not show a significant gradient ( < 5 K).Over Sm203 catalyst the reaction mainly occurred in the region ranging from zero to 11 mm from the front of the bed. Within this region, where the methane, oxygen and catalyst coexist, the methane was oxidized by oxygen species of the catalyst. Behind this reaction zone, no further conversion of methane was observed. In order to distinguish this part of the bed from the entire catalyst bed, we specified this region as the reaction zone of the bed. Within the reaction zone, the methane conversion increased linearly with bed height to a maximum at the end of the reaction zone and then remained constant. The linear increase of methane conversion as a function the bed height reveals that methane was oxidized with a constant reaction rate within this region. If the concentration of the active site on Sm,O, catalyst surface
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Catalysis A: General 115 (1994) 243-256
25 16
6
0
0 5
0
10
15
20
Bed Hciglu (nm) Fig. 2. The methane conversion and the concentration distributions of O2 and CO, along the Sm,Og catalyst bed at 873 K.
is assumed to be [ Oact] , methane is oxidized over these sites to form methyl radical and surface hydroxyl: CT& + [ Oact] -
Kl
CH3 * + [HO]
(I)
The reaction rate of the above reaction in fixed bed reactor can be expressed as
[I61 R=
VdX
-
dh
=
& pat400
(4)
where dXldh represents the reaction rate and corresponds to the slope of the methane conversion curves as a function of the bed height, PC+, is the concentration of methane in the flowing gas, r3o is the effective concentration of surface active oxygen [ Oact] when the catalyst is exposed to methane. The parameter K, 0, referred to the activity of the catalyst when it is working in the reactor. The variation in this parameter along the reaction zone can be calculated from the concentration distribution. As shown in Fig. 3, the activities of the working catalyst along the whole bed remain essentially constant. A comparison of the catalyst activity or the reaction rate with the concentration of oxygen and carbon dioxide along the catalyst bed (Fig. 2 and 3) indicates that 6Jodoes not vary with
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249
the change of gas concentration of oxygen and carbon dioxide. This result is consistent with those found previously for the Sm/CaO system [ 121, where the author found that the rate of methane conversion was also constant, by raising the reaction temperature to 1073 K to desorb carbon dioxide from the Sm/CaO surface. It is also expected on the basis of Langmuir adsorption, that the chemisorption of oxygen over Sm20, catalyst is saturated at low partial pressures ( <0.6 Pa) of oxygen. Results of Amorebieta and Colussi [ 17,181 gave support to this assumption, since the oxygen chemisorption reached saturation with less than 100 mPa of oxygen over Li/MgO [ 171 or Sm*O, [ 181 catalysts. When the reaction was carried out at 973 K over CaO catalyst, the conversion rate of methane and the oxygen concentrations are shown in Fig. 4. At the interface of the a-A1203 and the catalyst, the rate of methane conversion increases sharply, suggesting that the catalyst plays an important role in the conversion of methane. In the reaction zone (O-l 1.5 mm), the oxygen concentration decreases rapidly from the feed concentration of 11% to zero. The reaction rate dropped suddenly at a bed height of 11.5 mm, just at the point where oxygen was completely consumed. It is significant to note that as the gaseous oxygen reaches zero concentration at the end of the reaction zone, the rate of methane oxidation is still maintained at a high level (0.01 mm- ‘) . The same result has also been observed at 1073 K, while a higher reaction rate for methane conversion of 0.05 mm- ’ was observed. This sharp decrease in reaction rate indicates that neither the steady-state adsorption of oxygen nor a Langmuir-Hinshelwood mechanism involving unsaturated adsorption
l.O I t 0.8
.
0’
.
.
.
.
.
l
4k
3
‘3 0.6 e 1 = a v, 0.4 A
0.2
0.0
-
-
-
2
0
2
4
6 Bed
8
10
12
Heigd (mm)
Fig. 3. The catalyst activity ICC&,distribution along the Sm,O, catalyst bed at 873 K.
250
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Catalysis A: General 115 (1994) 243-256
10
5
0 -6 -4 -2
0
2
4
6
8
10 12 14 16 18 20 22
Bed Height(xm) Fig. 4. The methane conversion at 973 K.
rate distribution
and the. O2 concentration
distribution along the CaO catalyst bed
of oxygen lead to rate expressions compatible with the observed kinetic behavior. From a consideration of the constant activity of Smz03 in the reaction zone and assuming a Langmuir-Hinshelwood mechanism for oxygen adsorption, it may be concluded that oxygen adsorption is saturated under our reaction conditions, either in dissociative form or in molecular form. The concentration of oxygen in the reaction zone is high enough to result saturation of the adsorption at the operating temperature. This results implies that the concentration of active sites on the catalyst surface is not influenced by the variation in oxygen concentration. The decrease in methane conversion rate in the reaction zone over CaO catalyst may be a consequence of a different mechanism. This topic will be discussed below. 3.2. The role of carbon dioxide When the reaction was performed over CaO catalyst at the same temperature of 873 K as that used for the SmzO, catalyst, the CaO was found to be less active than Sm203. The rate of methane conversion and carbon dioxide distribution along the catalyst bed are shown in Fig. 5. The total conversion of methane was about 1.2%. The main products were the total oxidation products carbon dioxide (selectivity <94%) and carbon monoxide. The catalyst essentially behaved as an oxidation
C. Tang et al. /Applied
0.30
Catalysis A: General 115 (1994) 243-256
251 1.0
-
0.25 A -7
b
--_)
0.8
.
% “a 0.20 = $ d e 0.15 .B E 8 0
;
0.10
1 3
0.2
0.05
0.0
0.00 -2
0
2
4
6
8
10
12
14
16
Bed Height @nn) Fig. 5. The methane conversion rate distributionand the CO2concentrationdistributionalong the CaO catalyst bed at 873 K.
catalyst. Trace amounts of hydrocarbon products (C,H, and &I&) were also detected. Their concentrations along the catalyst bed varied in a similar way to that of carbon monoxide, but were estimated to be less than one tenth of that of carbon monoxide. At the beginning of the reaction zone, the rate of methane conversion increased rapidly to a maximum and then deciined along with the bed height. At a bed height of 9 mm, the reaction rate was essentially zero, while 11% oxygen and 38.6% methane were still present in the gas phase. This results indicates that the catalyst particles located behind the reaction zone are strongly deactivated by carbon dioxide, since large amounts of the reactants, methane and oxygen, co-existed with the catalyst. It is obvious that most of the active sites are occupied by some species which inhibit the adsorption and reaction of methane. In order to find which species play such an important role, temperature-programmed desorption (TPD) of carbon dioxide was performed after the reaction. Before TPD was performed, the reactor was flushed with helium to remove species that could easily desorb. The reactor was then cooled to ambient temperature. It can be seen from the TPD profile of Fig. 6 that a large amount of carbon dioxide was detected, desorption of carbon dioxide commenced at 873 K and reached a maximum at 973 K (shown in Fig. 6). Fig. 6 indicates that a large quantity of carbon dioxide was adsorbed on the catalyst surface. Adsorption of carbon dioxide blocks the surface sites and inhibits the methane conversion. This provides a clear
252
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4
600
1000
800 T-i=-=
1100
1300
(K)
Fig. 6. TPD spectrum of CO1 over CaO catalyst after OCM reaction at 873 K.
explanation for the decrease in the methane conversion rate of CaO. Initially there was no carbon dioxide in the feed gas, so the surface active sites are all exposed to methane and the reaction rate reached its maximum. As carbon dioxide was formed by the reaction, it covered part of the actives sites and reduced the activity of the catalyst. At a certain height of the bed, the concentration of carbon dioxide was high enough to cover all the surface active sites, and prevent the methane molecule to contact the active oxygen sites, thus resulting in the complete deactivation of the catalyst for methane oxidation. In Fig. 3, it is shown that the rate of conversion of methane remained constant over the whole bed of Sm20, catalyst even when carbon dioxide was present. TPD of carbon dioxide over this catalyst does not show any desorption. This results partly agrees with the results of Peil et al. [ 131 and Dingjun et al. [ 11I. They found that carbon dioxide has less inhibition effect on methane conversion over Sm203 than that over Li/MgO catalyst. Dingjun et al. [ 1l] recently reported that CH3 radical formation over Sm203 is affected much less by carbon dioxide than over Li/MgO catalyst. The carbon dioxide inhibition over CaO and other basic catalysts as MgO, SrO, CaO has also been observed by many groups [ 14,151, however, none observed the significant effect of produced carbon dioxide on active oxygen sites. 3.3. The rate equation The results of concentration distribution measurements provide a basis for quantitatively relating the activity of catalyst with the real concentration of carbon
C. Tang et al. /Applied
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253
dioxide, according to the reaction scheme. As mentioned above, the active sites were partly covered by produced carbon dioxide, COZ + Oact + [ COJ]
(II)
Under our special experimental conditions, the carbon dioxide adsorption is in equilibrium with the gas-phase concentration at a specific height of the catalyst bed. If the adsorption can be described by a Langmuir relationship, the coverage of COZ ( 0,,) on active sites follows:
632 =
K’PCO2
(3
1 + K’Pco2
where, Pcoz is the partial pressure of CO2 at bed height h, K’ is the equilibrium coefficient. In the previous discussion, we found that the oxygen concentration in the gas phase does not affect the reaction rate. As a result, the concentration of surface active sites in the reaction bed without carbon dioxide coverage can be assigned a value of unity. Therefore the effective concentration of active sites can be expressed as
00=1-ec,=
1 1 + K’Pco2
Finally, the methane conversion rate can be expressed as: R=
WC
-
dh
=
KIPcH., 1+ K’P,,
or
In the reaction, K1, K’, and V were constants, PcHI, PC&, and dX/dh can be obtained from the concentration distribution. It should be noticed that PC&, Pco, referred to the actual concentration over the working catalyst surface, which differs from that obtained from GC analysis in the effluent gas from the reactor. Since conversions at 873 K were very low, results at 973 K and 1073 K only are exhibited in Fig. 7. The results showed that PcH4/ ( dX/dh) varied linearly with carbon dioxide gas concentration. This correlation confirms the validity of Eq. (7) for quantitative expression of the inhibition of carbon dioxide. The adsorption equilibrium coefficient K’ can be obtained from the slopes of the lines in Fig. 7. The values of K’ at 973 K and 1073 K are equal to 0.487 Kpa-’ and 0.306 KPa-‘, respectively. At 973 K, the carbon dioxide concentration at the reaction zone is 5.15 KPa, the effective concentration of surface active oxygen 0, is 0.29. This value implies that 71% of the active sites are covered by produced carbon dioxide. At 1073 K, only 40% of the active sites were exposed to methane at the end of reaction zone. The
C. Tang et al. /Applied Catalysis A: General I15 (1994) 243-256
254
1
0
I,,, 0
(
,,,,,,,,,,,(,,i
1
!
,/,,,,,,,.
2
3
‘I
5
6
Pm2 SW
Fig. 7. PC,, / (dXldh) versus CO2pressureP c- in the reactionzone of the CaO catalyst.
result indicate that the produced carbon dioxide plays a very important role in methane oxidation over CaO catalyst or a similar basic oxide. To study the OCM reaction over this type of catalyst, the role of carbon dioxide must be carefully considered. It has been pointed out that carbon dioxide plays a role as a poison for the oxidative coupling reaction and increases the activation energy for the conversion of methane [ g-l 11. Apart from Roos et al. [ 71 who have considered carbon dioxide as a competitive adsorbate with methane, the inhibiting effect of carbon dioxide has not been considered in kinetic equations for OCM [ 19-231. Based on the results of normal catalytic fixed-bed reactors, many researchers showed that the reaction rate was related to the oxygen concentration in the feed stock, which seems in conflict with our results. It is worth to note that these results considered the catalyst bed as a whole. In our experiments, however, the oxygen concentration and the methane conversion are related directly to different layers of the working catalyst bed. These results have been obtained in one reaction run rather than in many experiments. This method will cause much less temperature variance on the working catalytic surface. In order to avoid severe temperature gradients within the catalyst bed and maintain the oxygen conversion rate below 20%, limited amounts of catalyst loading with large volumes of diluted inert mate-
C. Tang et al. /Applied Catalysis A: Geneml I1 5 (I 994) 243-256
255
rial are normally employed at high temperature ( > 973 K).Since a large amount of reaction heat was produced in a small amount of catalyst with a low surface area, the surface temperature of the catalyst is expected to vary extensively in different oxygen conversions. In order to prevent local heating of the catalyst, a large amount of diluent gas in the feed has usually been used. During these experiments, the oxygen concentration is rather low in comparison to our reaction conditions, thus, an unsaturated adsorption of oxygen is expected. As the reaction is run at low temperature, inhibition of the reaction rate by carbon dioxide would be very serious. In some cases, such as CaO, the catalyst behind the reaction zone exhibits no activity for methane oxidation. As a result, the conversion rate per gram of catalyst must be used carefully to express the reaction rate.
4. Conclusion The concentration profile of products and reactants as a function of distance along the catalyst bed has been studied for the OCM reaction over CaO and Sm203 catalysts by means of a specially designed system. It is found that the reaction mainly occurred in the reaction zone of the bed where the oxygen still exists in the gas phase. The concentration of active sites of CaO catalyst is not affected by the concentration of oxygen in the gas phase of the catalyst bed. The active sites of CaO catalyst can be deactivated by the carbon dioxide produced, and this is the main reason for the decrease in the methane conversion in the reaction zone. At low temperature (873 K), the catalyst can be completely deactivated when the partial pressure of produced carbon dioxide was greater than 0.8 kPa. Coverage of carbon dioxide to active oxygen sites can be quantitatively expressed by the Langmuir relationship. In contrast to the CaO catalyst, Sm203 catalyst does not show any inhibition effect by carbon dioxide under the same reaction conditions. All the results showed that the method for measuring the concentration distribution of reactant and product along catalyst bed was successful and effective in the study of the performance of OCM catalysts.
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