Effects of the addition of zeolites on ruthenium catalysts in carbon monoxide hydrogenation

Applied Catalysis, 49 (1989) 45-53 Elsevier Science Publishers B.V., Amsterdam -

45 Printed in The Netherlands

Effects of the Addition of Zeolites on Ruthenium Catalysts in Carbon Monoxide Hydrogenation YU-WEN CHEN* and WEI-JYE WANG Department of Chemical Engineering, National Central University, Chung-Li 32054, Taiwan (Republic of China) (Received 15 February 1988, revised manuscript received 22 November 1988)

ABSTRACT The synthesis of hydrocarbons from synthesis gas was studied over a series of zeolites (HX, HY and CuX) mixed with Ru/TiOP catalysts at 1 atm, 473-573 K, H,:CO = 1 and GHSV = 1800 h-l. A Ru/TiO,-Si02 catalyst was used for comparision. It was found that the acidity of the zeolites could modify the catalytic properties by ( 1) increasing CO conversion, (2 ) increasing C, hydrocarbon formation, (3) increasing the selectivites for olefins and (4) increasing isobutene formation. The results can be interpreted with a bifunctional mechanism of the catalysts. In the addition, the rate-determining step is shifted from the reactions taking place over the Fischer-Tropsch catalysts to diffusion of C, hydrocarbons into the zeolites.

INTRODUCTION

The major disadvantage of classical Fischer-Tropsch (FT) synthesis is that it produces a broad range of hydrocarbons from methane to waxes. Unsaturated hydrocarbons, especially those with two to four carbon atoms, constitute a large proportion of the feedstocks for the petrochemical industry. To date, the main problem has been that the FT catalysts which produce lower olefins do not give significant yields at high conversion levels [ 11. The development of highly selective ruthenium catalysts for the FT reaction has been studied extensively. Evidence has shown that molecular sieve zeolites can influence the ruthenium metal particle size and also its activity and selectivity [l-9]. Recently, research into new routes for improving product selectivity, especially for light olefins, using chemically modified catalytic metals have been emphasized. Alkali promoter and SMSI (strong metal-support interaction) supports have been used successfully [ 10-121. Modification of FT catalysts by mixing with zeolites is another means of enhancing the selectivity with respect to the desired products [ 13-221. Such a modification includes the use of catalytic systems in which two distinct types of active catalysts are intimately combined to provide the possibility of sequential reactions and in-

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46

creased selectivity with respect to certain products. Caeser et al. [ 131 found a synergistic effect of ZSM-5 zeolites in increasing selectivity with respect to aromatic hydrocarbons in the hydrogenation of carbon monoxide over an iron catalyst by using a mechanical mixture of these two components. In subsequent research, a Ru/A1203 catalyst was mixed with ZSM-5 zeolite and a similar synergistic effect in promoting aromatic hydrocarbon formation was observed [ 141. Obviously, zeolites can provide a dual catalytic function in a composite system to facilitate a significant shift in the normal product distribution [l-5]. Therefore, a reaction route for modifying hydrocarbon selectivites in the FT synthesis would be based on the interception of gaseous olefinic hydrocarbons before they can re-enter the surface chain growth reaction. The zeolite can be combined with the metal catalyst in various ways. In the one-step process, either the metal is loaded into the zeolite by impregnation or ion,exchange, or the two catalyst components are physically mixed. The resulting catalyst system is used in a single reactor. In t-hetwo-step process, two reactors are used in series, with the first reactor containing the metal catalyst and the second containing the zeolite catalyst. In this present study, an investigation was carried out on the FT synthesis over mixed catalysts consisting of Ru/TiO, and a series of zeolites. Ru/TiO, is an excellent catalyst for producing olefinic hydrocarbons [lo]. Zeolite catalysts active for the isomerization of a-olefins and for the catalytic cracking of hydrocarbons, such as HX, HY and CuX, were employed, The main idea is that the secondary reaction of olefinic hydrocarbons over these zeolites would provide a modification of the normal product distribution of the primary ruthenium catalyst. EXPERIMENTAL

The Ru/TiO, catalyst was prepared by incipient-wetness impregnation of 4 wt.-% of ruthenium on a titania support (Degussa P-25) with an aqueous solution of RuCl,-3Hz0 (obtained from Strem Chemicals) After impregnation, the samples were dried in air at 350 K for 50 h. The composite catalysts were prepared by grinding together a Ru/TiO, catalyst (C ) and a zeolite (Z) in an excess volumetric ratio (Z:C = 6) and screening to 150-200 mesh. The excess amount of zeolites could shorten the transient period of the reaction. It has been reported [ 161that the induction period is fairly long for FT-zeolite catalysts to reach a steady state. The duration of induction period is greatly shortened by using a high Z:C ratio. The zeolites employed were HY, HX and CuX. The Ru/Ti02 samples were taken from the same batch to prevent variations in dispersion. The X-ray line broadening analysis indicated that the ruthenium particle size is less than 40 A. The HX and HY zeolites were obtained from NH,+ ion-exchanged with NaX and NaY (Strem Chemicals ) followed by calcination at 773 K for 2 h in

47

air to remove NH3. The CuX zeolite was prepared by ion exchange of NaX with 0.2 A4 copper (II ) nitrate solution at pH 5 at room temperature. After ion exchange, the sample was washed with deionized water at 300 K followed by calcination in air at 673 K for 6 h. The degree of ion exchange (60% Na+ replaced by Cu2+ ) was determined by analysis of the copper solution before and after the exchange process. The X-ray powder patterns indicated that the zeolites were X and Y types with good crystallinity. A study of the catalyst activity and selectivity was conducted under typical FT reaction conditions, viz., 1 atm, 473-573 K, GHSV 1800 h-’ and H,:CO = 1. The reaction tests were carrier out in a quartz U-tube microreactor fitted with a thermowell. The catalyst bed temperature was controlled to within ? 1 K by means of a PID electronic controller. The catalyst was pretreated in situ under a flow of hydrogen at 693 K for 16 h. The catalyst temperature was lowered to the reaction temperature before introducing the feed gas. All the reactant gases were of high purity (99.9%; Matheson) and were further purified by passing them through Drierite and molecular sieve traps to remove water and hydrocarbon impurites. The hydrocarbon product distribution was determined by using an on-line gas chromatograph (HP 5890A) with a Porapak QS column and a 15% Apiezon-Chromosorb W AW column in series. Two detectors, Thermal-conductivity and flame ionization, were used in series. However, not all the C, components were well separated by this column. Therefore, a Carbopack C-0.19% picric acid column was used for complete separation of all the C!, components. The product gases were transferred to the gas sampling valves via heated tubes at 430 K. All the peak areas were measured by Hewlett-Packard electronic integrators. The calibration was carried out using various calibration mixtures. The carbon monoxide conversions obtained were in the range l-15% and the reproducibilities were less than 3%. The size of the reactor (3 cm~0.4 cm I.D.) was chosen on the basis of the following general criteria. First, the reactor has to be long enough that the influence of axial dispersion on steady-state conversion can be neglected. It has been shown that in steady-state isothermal operation at Reynolds numbers > 20, the axial dispersion effects can be neglected so long as the length of the reactor is greater than 50 particle diameters [ 231. In non-isothermal operation, the length-to-particle diameter ratio has to be increased three-fold in order to compensate for the three-fold decrease in the axial thermal Peclect number. Second, the diameter of the reactor has to be at least ten times the particle diameter so that the influence of radial dispersion and the wall can be minimized. This is based on the fact that the radial Peclect number is about 5-6 times greater than that of axial dispersion, which means that radial profiles are less easily damped by dispersion than the axial profiles. As a result, the ratio for the reactor diameter is 5-6 times smaller than that for the reactor length,

48 RESULTS AND DISCUSSION

The rate of a heterogeneous catalytic reaction is determined by a combination of serial and parallel rate steps. The rate steps are mainly composed of bulk diffusion, pore diffusion and intrinsic kinetics. It has been found in this study that changing the gas flow-rate, while keeping the GHSV constant, has little effect on the conversion efficiency. Hence the factor of bulk diffusion can be neglected. The rate of pore diffusion can be affected by pore structure and particle size. Pore structure can be characterized by the void fraction and tortuosity, both of which influence the effective diffusivity. The particle size, on the other hand, affects the effectiveness factor, a rating factor for the diffusion effect through a Thiele modulus [ 231. Increasing the catalyst particle size will increase the Thiele modulus which, in turn, decreases the effectiveness factor. In this study, a very small particle size (150-200 mesh) was used, such that the effectiveness factor became unity and no pore diffusion effect on conversion efficiency should be observed. However, it is extremely difficult to overcome the diffusional effect of C, hydrocarbons in the zeolite cage, as will be described later. In Table 1, typical catalytic activities of the SiOz-diluted Ru/TiO, at a reaction temperature of 553 K. The catalyst activity was determined from the reaction rate of overall carbon monoxide conversion. The results with SiOZdiluted Ru/TiO, were used for comparison in order to isolate the physical effects of the second component on the properties of mixed system [ 241. The results indicated that the Ru/TiOz-zeolite catalysts were more active than the SiO,-diluted catalysts. This is attributed to a multi-functional reaction, as suggested in the literature [ 15-221. The rate of C, hydrocarbon formation is also compared, as shown in Table 1. The mixed catalyst containing either HX and CuX were at least 60% more active in overall C, formation than the SiO,-diluted catalyst, probably as a TABLE 1 Catalytic activities of Ru/TiO,-zeolite Catalyst Ru/TiO,-SiO, Ru/TiO,-HX Ru/Ti02-HY Ru/TiO,-CuX

catalyst

- f-co’ (,umol g-’ s-‘)

E actb (kcal mol-‘)

bc (pmoig~ls-1

1.34 1.82 1.40 1.62

17.0 10.4 7.2 9.0

0.067 0.153 0.126 0.130

” Reaction conditions: 1 atm, 553 K, H,: CO = 1, GHSV = 1800 hh’. hActivation energy of CO disappearance. ’ Rate of formation of C, hydrocarbons. ’ Selectivity of C, hydrocarbons.

)

s d (Zt.-X) 5.0 8.4 9.0 8.0

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result of non-trivial secondary conversion of C, olefins, which may also have accounted for the increase in their activity in the overall CO conversion [ 211. The lower rate increase with the HY system in C, formation was possibly due to the cracking ability of HY, which produced coke during the reaction. The activation energy of the reactions over the zeolite-containing catalysts is about 7-10 kcal mol-‘, which is substantially lower than the 17 kcal mol-’ for the SiO,-diluted catalyst (Table 1 and Fig. 1). Similar observations have been made with composite iron/zeolite catalysts for FT synthesis [ 131; an activation energy of 5 kcal mol-’ was found for mixed Fe/ZSM-5, compared with about 15 kcal mol-1 for the primary iron catalyst. In a non-trivial reaction, it is expected that a shift in activation energy would be observed [ 251 as the overall mechanism has shifted from an FT mechanism to a sequential reaction mechanism involving active olefin transformation. Hydrocarbon product distributions over these catalysts are shown in Figs. 2-4. HX and CuX mixed catalysts had similar CH4 selectivities to the primary Ru/TiOz-SiOz but modified light olefin production toward secondary products. The higher light olefin formation from the HX and CuX mixed catalysts was attributed to the secondary reaction on the acidic sites of the zeolite surface. The increase in the relative methane fraction cannot result from secondary cracking reactions, as cracking of long-chain olefins through a carbonium ion intermediate does not produce methane [ 261. It has been hypothesized that pore condensation in zeolites may also be responsible for the limitation of long-chain hydrocarbons, We tried to extract high-molecular-weight hydrocarbons to check for this, but did not find them. In an additional study, the HX zeolite was placed at the tail end of the reactor bed in a separate layer. Under FT reactions, this two-layer catalyst was found to have the same activity and product distribution as a single layer of Ru/TiO, catalyst. The results clearly indicate that the pore condensation of hydrocar-

9

1 l/K)

Fig. 1. Arrhenius plot. Reaction conditions: 1 atm, H,:CO= 1, GHSV= 1800 h-l.

51

from the Weisz-Prater

criterion

[ 251.

Diffusion limitation

is more apparent

in zeolites. For the diffusion of C, hydrocarbons in zeolites, the diffusivities have been found to be in the range 1. lOW-1. lo-’ cm2 s-l [ 27-291. This also explains the decrease in activation energies for the zeolite-mixed catalysts, as the apparent activation energy would be reduced to half of its original value for a diffusion-limited reaction. Varma et al. [ 211 reported that the rate-determining step is shifted from the reactions taking place over the FT catalyst to those over the zeolite catalyst. If this is so one would expect the diffusion of C, hydrocarbons into zeolite to be the rate-determining step in our catalyst system. The effects of temperature on product selectivities are also shown in Figs. 2-4. The selectivities for methane are, as expected, increased with increasing temperature. It is known [7] that the Gibbs free energies of reaction for the formation of high-molecular-weight hydrocarbons in the FT synthesis are more negative and decrease more rapidly with decreasing temperature than that for methane under FT conditions (493-593 K). The selectivities for C,-C, olefins are decreased with increasing temperature. This behaviour is normal as the olefinic hydrocarbons are less stable under the reaction conditions and can be hydrogenated successfully to alkanes. It should be noted that the formation of alkanes is thermodynamically favoured at FT temperatures. As the relative differences between the Gibbs free energies of the alkanes and olefins decrease as temperature decreases, it has been suggested that this should result in an increase in the olefin fraction as the reaction temperature is decreased [ 331. However, such an explanation cannot possibly explain the formation of olefin fractions that exceed those of alkanes. Thegreatest variance in the olefin fraction probably results from differences in the various kinetic steps. It has been suggested that olefins are primary products in FT reactions and their fraction decreases with increasing temperature owing to an increase in hydrogenation activity [ 71. A significant part of the C4 fraction was in the form of isobutene for zeolitemixed catalysts. Owing to the bifunctional properties of the catalysts, it is not surprising that branched-chain hydrocarbons would be found in the product stream: isobutene l-butene

Ru

isobutane

Llite

n-butane Various workers [3,30] have found that the acid strength of the OH group appeared to be important in affecting the fraction of branched hydrocarbons obtained. Our results are in accord with this. In a previous study [3] it was

52

found that on zeolite-supported ruthenium catalysts the major component of the C, fraction is isobutane. However, in this study, we found large amounts of isobutene rather than isobutane. The results can be attributed to the low hydrogenation ability of SMSI Ru/TiO, catalysts [ 10,111. Lee and Ihm also found the same phenomenon in studies over ion-exchanged Co/Y catalysts [31,32]. The role of mass transfer in bifunctional catalysis has been discussed in detail by Weisz [ 251. Pore diffusion is likely to be particularly important for the production of isobutane in the present system. Weisz [25] derived a general criterion to check for the absence of mass-transport limitations, and many workers [ 3,13,33-351 have used this criterion with zeolite catalysts. Therefore, Weisz’s criterion [ 251 was applied to check for the presence of absence of masstransfer limitations in this study. The criterion for the absence of mass transport limitations is

where @ is Weisz’s number, dn/dt is the reaction rate for isobutane formation per unit volume of catalyst, [B],, is the concentration of isobutene at equilibrium, R is the radius of a zeolite palticle and D is the diffusivity of isobutene. From the observed conversion, the rate of isobutane formation per unit volume of the catalyst was calculated to be 1.10-l mol s-l cm-‘. The concentration of isobutene at equilibrium is calculated to be about 1.2 *lo-’ mol cm-3 at the reaction temperature of 523 K [ 361. The effective diffusivity of isobutene in zeolite has been reported to be 1.0. 1O-6 cm* s- ’ at 523 K [ 271 and the radii of the zeolite particles are about 0.01 cm. Substituting these values into Weisz’s criterion, we obtained 0~1, which indicates that the reaction was substantially constrained by diffusion limitations of 1-butene in zeolite cages. It should be noted that the isobutene-to-n-butene ratio is considerably in excess of the equilibrium value [ 361. This reinforces the proposition that isobutene is formed mainly via the bifunctional mechanism and the diffusion limitation of 1-butene in zeolites. If isobutene is the primary product on ruthenium metal sites, its concentration would be restricted by equilibrium. ACKNOWLEDGEMENT

This research was sponsored by the National Science Council of the Republic of China under grant number NSC77-0402-E008-02.

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