Conversion of syngas to liquid hydrocarbons over a two-component (Cr2O3–ZnO and ZSM-5 zeolite) catalyst:

Chemical Engineering Science 55 (2000) 1845}1855

Conversion of syngas to liquid hydrocarbons over a two-component (Cr O }ZnO and ZSM-5 zeolite) catalyst: 2 3 Kinetic modelling and catalyst deactivation Javier Eren8 a!,*, Jose M. Arandes!, Javier Bilbao!, Ana G. Gayubo!, Hugo I. De Lasa" !Departamento de Ingeniern& a Qun& mica, Universidad del Pan& s Vasco, Apartado 644, 48080 Bilbao, Spain "Chemical Reactor Engineering Centre, University of Western Ontario, London, Ont. N6A 5B9, Canada Received 4 December 1997; accepted 8 July 1999

Abstract The present study describes the kinetics of syngas transformation into liquid hydrocarbons (boiling point in the gasoline range) using as catalyst a mixture of a metallic component, Cr O }ZnO, and of an acidic component, ZSM-5 zeolite. Experimental results 2 3 were obtained in an isothermal "xed-bed integral reactor. The validity of several kinetic models, available for methanol synthesis, is analysed and modi"cations are proposed. These changes involve a rate equation with a CO concentration-dependent term. Catalyst 2 deactivation is also evaluated and the e!ect of the operating conditions on coke deposition is established. Moreover, the rate of CO conversion and the change of catalytic activity with time-on-stream were described using a kinetic model showing a weak in#uence of temperature. ( 2000 Elsevier Science Ltd. All rights reserved. Keywords: Syngas conversion; Cr O }ZnO/ZSM-5 catalyst; Cr O }ZnO catalyst; ZSM-5 zeolite; Catalyst deactivation; Synthesis of hydrocarbons 2 3 2 3

1. Introduction The transformation of syngas into liquid hydrocarbons has deserved signi"cant attention since bifunctional catalysts were proposed by Chang, Lang and Silvestri (1978). This approach provides an alternate route for exploitation of natural gas, coal and biomass and for the production of automotive fuels. The catalyst, composed of metallic and acidic components, allows for the synthesis of hydrocarbons in a single reactor with the metallic component catalyzing the synthesis of methanol and the acidic component catalyzing the selective transformation of methanol into hydrocarbons. A number of valuable contributions con"rmed the potential value of this approach (Chang, Lang & Bell, 1980; Inui & Takegami, 1982a,b; Costa-Novella, Calleja & Ballesteros, 1984; Yashima, Yoshimura, Wakuskima & Namba, 1984; de Lasa, Ravella, Rost & Mahay, 1989; Simard, 1991; Simard, Mahay, Ravella, Jean & de Lasa, 1991; Simard, SedraH n, SepuH lveda, FmH goli & de Lasa, 1995; Comelli & FmH goli, 1993). In particular, it was observed that the coexistence of both steps, methanol synthesis and

* Corresponding author.

methanol conversion, removes the thermodynamic equilibrium limitations for methanol synthesis, favouring the overall syngas transformation (Marschner & Moeller, 1983). Eren8 a, Arandes, Bilbao, Aguayo and de Lasa (1998a) considered a number of aspects relevant to these catalyst with two components, showing the good behaviour of a Cr O }ZnO/ZSM-5 catalyst (Chang et al., 1978,1980; 2 3 Costa-Novella et al., 1984; Yashima et al., 1984; de Lasa et al., 1989; Simard, 1991; Simard et al., 1991,1995). This was achieved comparing the performance of this catalyst with others, also prepared by physical mixing of an acidic component (ZSM-5 zeolite) with metallic components conventionally used for methanol synthesis. In this respect, Eren8 a et al. (1998a) established that, for optimal selectivity towards gasoline range hydrocarbons, the catalyst has to have an Cr/Zn atomic ratio of 2.0 in the metallic component and a Si/Al ratio of 154 in the acidic component. This corresponds to a metallic component/acidic component mass ratio of 2.0. Furthermore, other recent studies (Yashima et al., 1984; Simard et al., 1991; Eren8 a et al., 1998a, Eren8 a, Arandes, Bilbao, Olazar & de Lasa, 1998b), using the aforementioned catalyst consider the e!ect of various operating conditions. It was observed that, in most cases,

0009-2509/00/$ - see front matter ( 2000 Elsevier Science Ltd. All rights reserved. PII: S 0 0 0 9 - 2 5 0 9 ( 9 9 ) 0 0 4 3 8 - 8

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Nomenclature ARE a, a 0 C c C i d F k k d K %2 K i ¸ m, n N p i p i0 P

error objective function (average relative error), Eq. (1) catalyst activity at t and zero time-on-stream, referred to reaction rate, Eq. (24) coke content, referred to the catalyst concentration of i component, mol m~3 deactivation order combined molar #ow of (CO#H ) react2 ants, mol h~1 kinetic constant deactivation kinetic constant, h~1 equilibrium constant for methanol synthesis chemical species constant as reported in Tables 2 and 3 reactor length, m exponents in Eqs. (12) and (22) number of experimental points partial pressure of i component, atm partial pressure of i component at the reactor inlet, atm pressure, atm

the concentrations of CO , H O and hydrocarbons in 2 2 the product stream remain essentially unchanged. For example, typical product weight fractions are: 70}75 wt% for CO , 22}28 wt% for hydrocarbons, with small 2 amounts of water ((5 wt%). This product distribution agrees well with the following reaction scheme: nCO#2nH HnCH OH, 2 3 nCH OHP(CH ) #nH O, 3 2n 2 nCO#nH OHnCO #nH . 2 2 2 In fact due to the presence of water in small amounts, the experimentally observed overall reaction stoichiometry can in practice be approximated as 2nCO#nH P(CH ) #nCO . 2 2n 2 This stoichiometry is somewhat unfavourable given the high levels of CO in the products. Nevertheless, and 2 in spite of this, there are still good prospects for this process, given the observed high reaction rates and the high conversions of CO towards gasoline with high octane index. Two major aspects are considered in this study: (a) establishing a kinetic model for quantifying the rate of syngas transformation; (b) the in#uence of the time-onstream on the kinetic model. On these matters, it should be mentioned that there is little information available and knowledge is needed to ascertain the industrial viability of this process. Given that the limiting reaction

r, r 0

reaction rate at t and zero time-on-stream, dX/d(=/F) R constant of gases, J mol~1 K~1 (atm l mol~1 K~1 in Eqs. (3) (6) (8) and (9) t time-on-stream, h t time-on-stream required for the catalyst to %2 reach stabilization, h ¹ temperature, K u ,u surface velocity of the gas at the inlet and 0 along the reactor, m h~1 = catalyst mass, g X,X experimental and calculated value of i comi i,#!-# ponent conversion z longitudinal coordinate along the reactor, m Greek letters a parameter de"ned in Eq. (7) e bed porosity e expansion parameter, Eq. (4) CO / fugacity of i component i o bulk density, g cm~3 m dimensionless coordinate along the reactor

step is methanol synthesis, with the reaction network favouring methanol transformation (de Lasa et al., 1989; Eren8 a et al., 1998a,b), it is interesting to analyse whether kinetic models for methanol synthesis with suitable modi"cations could adequately represent the overall synthesis gas conversion. While it was previously suggested that catalyst deactivation is relatively slow (Simard et al., 1991), the study of deactivation deserves more attention. Both pro"tability of this process on a large scale, as well as its comparison with other alternative routes for obtaining automotive fuels, depend mainly on the quanti"cation of catalyst deactivation. Knowledge of deactivation kinetics can also help to establish viable reaction}regeneration strategies.

2. Experimental A number of previous contributions report the di!erent preparation steps and the characterization of the metallic component, Cr O }ZnO (Eren8 a, 1996; Eren8 a 2 3 et al., 1998a), and of the ZSM-5 zeolite (Benito, Aguayo, Gayubo & Bilbao, 1994; Benito, 1995). The catalyst was prepared, in the present study, by physically mixing two components to reach a mass ratio of metallic component/acidic component of 2.0. Prior to the experiments, the catalyst was subjected to an equilibration treatment, by oxidation}reduction in the reactor

J. ErenJ a et al. / Chemical Engineering Science 55 (2000) 1845}1855

itself, as follows: (a) air was circulated for 30 min at 4003C; (b) a stream of He contacted the catalyst for 30 min at 4003C; (c) H was circulated for 2 h at 4003C. 2 The ratio of Bronsted/Lewis acidic sites in the ZSM-5 zeolite was 4.42. This ratio was determined following pyridine adsorption and from the ratio between the intensity of the FTIR adsorption bands at 1550 and 1455 cm~1 (Datka & Piwowarska, 1988). The total acidity of this catalyst, as measured by TPD, was 0.19 (mmol NH or of tert-butylamine) g~1. 3 The reaction equipment used (Autoclave Engineers BTRS Jr.) was limited to 100 atm and 6503C. The reactor was an isothermal "xed-bed (6.4 mm i.d., 152.4 mm length) with up to 5 cm3 of catalyst. Temperatures in the bed and on the internal wall of the reactor were measured with 0.5 mm chromel-alumel thermocouples with a $13C accuracy. These temperatures were controlled by means of a Eurotherm 847 controller, so that temperature di!erences between the bed and the wall of the reactor were always lower than 13C (Eren8 a, 1996). The on-line analysis of the reaction products (45 have been identi"ed) were done with a Hewlett-Packard 5890 Series II gas chromatograph, equipped with a thermal conductivity detector, TCD and a #ame ionization detector (FID). Three columns were used: (1) a semicapillary column (HP-1 Crosslinked Methyl Silicone), which separates the volatile (C }C ) and polar (methanol, 1 4 water and dimethyl ether) components from the remaining products; (2) a packed column (Porapak Q 80}100 mesh) separates the volatile components to be analysed by TCD; (3) a capillary column (PONA Crosslinked Methyl Silicone) which separates the heavy products to be analysed by FID. The optimal conditions for these products to be analysed were: initial temperature, 753C (7 min); ramp, 203C min; "nal temperature, 2403C (6 min). The collection and data analysis were done with Hewlett Packard software, HP 3365 Chemstation.

3. Results Two groups of experiments were developed as follows: (a) at constant temperature using increasing times-onstream; (b) with a temperature}time selected ramp at various times-on-stream. The latter technique reduced the experimental e!ort, as in each experimental run a signi"cant amount of kinetic data for di!erent temperatures were obtained. The experiments at constant temperature were carried out at the following conditions: (a) catalyst: Cr O }ZnO/ZSM-5 (metallic component/acidic com2 3 ponent ratio: 2.0; Cr/Zn atomic ratio: 2.0; Si/Al ratio:154; (b) mean particle size: 0.20 mm; (c) pressure: 30 atm; (d) temperature: 350, 375, 400 and 4253C; (e) space time: 7.23, 14.47, 28.93 and 62.22 (g of catalyst) h (mol(CO#H ))~1; (f) time-on-stream: between 10 min 2

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Fig. 1. Changes of CO conversion with time-on-stream, for di!erent temperatures. =/F"28.93 (g of catalyst) h (mol of (CO#H ))~1. 2

Table 1 Composition of the gasoline fraction (C ) at 4003C and for di!erent 5` space time values C distribution 5` wt %

Space time (g of catalyst) h (mol (CO#H ))~1 2 7.23 14.47 28.93 66.22

C 5 C 6 Benzene C 7 Toluene C 8 C aromatics 8 C aromatics 9 C aromatics 10 C aromatics 11` Total

30.8 25.4 1.1 13.7 3.8 6.0 14.5 3.8 0.7 0.1 100

24.0 19.9 1.0 11.3 3.1 5.0 23.0 9,8 2.7 0.2 100

19.7 11.0 0.5 6.9 2.2 3.3 26.8 22.0 7.4 0.2 100

15.2 5.0 0.2 2.6 5.2 0.4 25.1 35.2 10.9 0.2 100

and 22 h; (g) CO/H molar ratio: 0.5; (h) molar #ow of 2 reactants: 1.15 mmol(CO#H ) min~1. 2 Fig. 1 displays the conversion of CO with time-onstream for the space time 28.93 (g of catalyst) h (mol(CO#H ))~1. There was an initial period when 2 the CO conversion increased and this corresponded to the so-called non-equilibration period of the metallic component. At the end of this period the metallic component reached a stable oxidation}reduction state (equilibrium condition). Eren8 a et al. (1998b) observed that the duration of this stabilization period depends on the operating conditions, especially on the composition of the metallic component and on the CO/H ratio of the 2 incoming feed mixture. Table 1 shows the distribution of products in the fraction C for 4003C, once the stabiliz5` ation period has elapsed, for di!erent values of space time. Kinetic models were "tted using the pseudo-equilibrium condition for the maximum CO conversion. This

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integral reactors was adopted here. Discrimination and estimation of kinetic parameters involved minimization of the following average relative error objective function (Box, 1965): D/X )2 +N (DX !X i,#!-# i , ARE" i/1 i N

Fig. 2. Changes of CO conversion with time-on-stream for various W/F and a temperature}time ramp of 0.53C min~1.

state was attained in &2 h (t"t ) and was considered %2 as a reference time (t!t "0) for further modelling. %2 Catalyst performance was only modi"ed thereafter by slow catalyst deactivation. Given that catalyst deactivation is much slower than the period needed to reach equilibration, both periods * initial stabilization and deactivation * were studied independently. The second group of experiments was done in the following sequence: (a) an initial temperature of 3253C for 270 min; (b) a temperature ramp at 0.53C min~1 until 4253C, (c) a "nal temperature of 4253C for 120 min. Fig. 2 shows CO conversions for di!erent times-onstream when temperature was increased in a programmed manner. Each curve in Fig. 2 corresponds to one value of space time. As a result of the aforementioned catalyst stabilization an initial increase in CO conversion is also observed in Fig. 2. It will be noticed that CO conversions during the periods of increasing temperature are nearly identical to those obtained when operating at constant temperature (refer to maximum CO conversions in Fig. 1). Product distribution is very close as well and this even includes the distribution of hydrocarbons and of the components of the gasoline fraction. The similarity between results at constant temperature and at variable temperatures shows that, once the catalyst equilibrium state is reached at say 3253C, as temperature increases additional time-on-stream for subsequent catalyst equilibration is not required. This result becomes apparent in the range between 325 and 4253C, given that the oxidation-reduction state of the active catalyst components remains essentially unchanged.

(1)

where X is the calculated CO conversion obtained by i,#!-# numerically solving using orthogonal collocation the mass conservation equations for the i chemical species involved (Villadsen & Michelsen, 1978). Considering sound assumptions such as minimum axial dispersion, negligible interparticle and intraparticle di!usional limitations and also accounting for catalyst deactivation, there results LC LC Lu 1!e i "!u i !C ! o[!r (C , ¹)]a. i Lz 0 i Lt Lz e

(2)

Eq. (2) becomes Eq. (3) considering the following: (a) a dimensionless coordinate m"z/¸ is introduced; (a) temperature changes with time-on-stream in an isothermal reactor with temperature increasing following a ramp function; (c) the ideal gas law can be adopted: Lp p L¹ i! i Lt ¹ Lt u Lp p Lu 1!e i! i ! "! oR¹[!r (p , ¹)]a. 0 i ¸ Lm ¸ Lm L

(3)

Furthermore, for a chemical reaction with changes in volumetric #ow, the (p /¸ Lu/Lm) group can be expressed i as a function of (Lp /Lm) using the following two equai tions: u"u (1#e X ) 0 CO CO

(4)

and (1!X ) CO . p "p 0 CO CO (1#e X ) CO CO

(5)

Thus, considering Eqs. (4) and (5), Eq. (3) becomes: Lp p L¹ CO ! CO Lt ¹ Lt

A B

"! 4. Kinetic modelling of the syngas conversion

u Lp 1!e CO ! !a oR¹[!r (p , ¹)]a 0 i ¸ Lm e

(6)

with 4.1. Methodology for the kinetic study A similar methodology to the one proposed by Gayubo, Arandes, Aguayo, Olazar and Bilbao (1993a,b, 1994), Gayubo, Llorens, Cepeda, Olazar and Bilbao (1997a) for complex reactions in isothermal "xed-bed

p u e p (1#e ) CO . a" CO 0 CO0 CO0 ¸(p 0 #e p )2 CO CO CO

(7)

The term (p ¹/L¹/Lt) in Eq. (6) becomes zero when CO there is no variation of temperature with time-on-stream.

J. ErenJ a et al. / Chemical Engineering Science 55 (2000) 1845}1855

Thus, under these conditions Eq. (6) can be expressed as

A B

LX u LX CO "! !a CO Lt ¸ Lm (1!e) (1#e X )2 CO CO oR¹[!r (p , ¹)]a. # (8) 0 i p 0 (1#e ) e CO CO Moreover, the condition t"t and once the catalyst %2 reaches stabilization, the CO conversion along the bed can be calculated solving the mass conservation equation, Eq. (8), with the LX /Lt set to zero: CO LX CO " Lt [(1!e)/e][(1#e X )2/p 0 (1#e )]oR¹[!r (p , ¹)]a CO 0 i 0, CO CO CO ! (u/¸!a) (9) where a is the initial activity with a value of 1 for the 0 fresh catalyst. Calculation of the kinetic parameters was developed using reparameterization, thus reducing parameter crosscorrelation (Kittrell, Mezaki & Watson, 1965; Mezaki & Kitrell, 1967; Agarwal & Brisk, 1985). As stated, the kinetic modelling of syngas conversion at t"t was carried out by solving Eq. (9) using di!er%2 ent kinetic models. Results obtained by solving this equation were identical to those obtained by applying the integral method of data analysis to the corresponding kinetic model. Finally, a kinetic study of catalyst deactivation was done by solving Eq. (8) at di!erent times-onstream. 4.2. Modixcation of kinetic models for methanol synthesis After the pioneering work of Natta (1955) a signi"cant number of syngas transformation studies (with or without CO ) into methanol have been made (Bakemeier, 2 Laurer & Shroder, 1970; Schermuly & Luft, 1977; Klier, Chaticavanij, German & Simmons, 1982; Ostrovskii & Dyatlov, 1982; Matulewicz, De Keijser, Mol & Kapteijn, 1984; Seyfert & Luft, 1985; Shub, 1983; Agny & Takoudis, 1985; Skrzypek, Grzesik & Szopa, 1985; Skrzypek, Lachowska & Moroz, 1991; Kuczynski, Browne, Fontein & Westerterp, 1987; Graaf, Stamhuis & Beenackers, 1988; Kuechen & Ho!mann, 1993; Coteron & Hayhurst, 1994). Table 2 reports the values of the error objective function, Eq. (1), and of the kinetic parameters obtained by "tting the experimental results of this work to di!erent available models. It should be pointed out that under the experimental conditions used here the concentrations of both methanol and water in the product stream were negligible. Consequently, the kinetic models were "tted without taking into account terms corresponding to the concentration of methanol and water. Thus, the models

1849

set out in Table 1 are those corresponding to syngas transformation, without CO as reactant in the feed. The 2 aim of the study is not to evaluate the kinetic equations proposed for methanol synthesis, but rather to test the applicability of these equations in a more complex process, in which methanol synthesis is the "rst step and which, on the whole, is strongly in#uenced by the subsequent step of methanol transformation. The best "tting to the experimental results corresponds to the models of Schermuly & Luft (1977) and of Bakemeier et al. (1970), with values of the error objective functions of 1.37]10~3 and 1.45]10~3, respectively. A comparison between the experimental values of CO conversion and the values calculated from these models is shown in Figs. 3 and 4, respectively. It is speculated that the suitability of these kinetic models is due to the presence in the denominator of a term depending on the concentration of CO . In the case of the model of 2 Bakemeier et al. (1970), the adequate "tting is also due to the optimization of the exponent values for the partial pressures of CO and H . The model of Graaf et al. (1988) 2 also "ts the experimental results well (Fig. 5), as do the models of Schermuly & Luft (1977) and Bakemeier et al. (1970), when the presence of a CO concentration depen2 dent term in the denominator signi"cantly helps in this respect. Nevertheless, the predictions of the model of Graaf et al. (1988) deviate appreciably for high values of space time. The remaining kinetic models in Table 2, except the model of Seyfert and Luft (1984) do not include a term in [CO ] in the rate equation's denominator, and this may 2 explain the poor "tting and the high values of the error objective function obtained. For the model of Seyfert and Luft (1984), however, the inclusion in the denominator of an additional K/ /2 2 group is inadequate, as may be CO H observed when compared with the "tting obtained with the model of Schermuly & Luft (1977). Considering these "ndings and in order to improve the "tting of the kinetic models reported in the literature, di!erent modi"cations to those models are proposed here. Table 3 includes additional terms, dependent on the concentration of CO , for the models proposed by Natta 2 (1955), Matulewicz et al. (1984) and Kuczynski et al. (1987). An empirical model proposed by Agny and Takoudis (1985) and by Skrzypek et al. (1985) was also considered in which the values of the exponents of the concentrations of CO and H are adjustable. 2 Table 3 reports the ARE (see (1)), and the kinetic constants for best "tting. It can be concluded that the model of Natta (1955) with a term in [CO ], is the one 2 providing the best "t. This equation shows the lowest ARE"1.08]10~3. Fig. 6 illustrates the adequacy of the data "tting for the conditions at t"t . %2 As reported in Table 3, the modi"ed model of Matulewicz et al. (1984) does not improve the "tting of the original model. Nevertheless, the addition of a term

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Table 2 Kinetic constants and error objective function for the di!erent kinetic models in the literature Kinetic model

k(/ /2 !/ /K ) CO H2 CH3 OH %2 Natta (1955) r" (1#K / #K / #K / )3 CO CO H2 H2 CH3 OH CH3 OH

(10)

Error objective function, Eq. (1)

Kinetic constants

0.121

k"0.112 exp(!9580/R¹) K "0.327]10~5 CO exp(32 920/R¹) K "0.106 exp(10 830/R¹) H2 k"0.748]10~1 exp(!12 080/R¹) K "0.110]10~4 CO exp(28 330/R¹) K"0.890]10~1 exp (11 250/R¹)

k(/ /2 !/ /K ) CH3 OH %2 CO H2 Kuczynski et al. (1987) r" (1#K / #K/ /2 ) CO CO CO H2

(11)

0.220

kpm pn (1!p /p p2 K ) CH3 OH CO H2 %2 Bakemeier et al. (1970) r" CO H2 (1#K /p ) CO2 H2

(12)

0.145]10~2

Schermuly and Luft (1977) k(/ /2 !/ /K ) CH3 OH %2 CO H2 r" (1#K / #K / #K / #K / )2 CO2 CO2 CO CO H2 H2 CH3 OH CH3 OH

0.137]10~2

(13)

k"3.590 exp(!42 920/R¹) K"2.952 exp(26 670/R¹) m"2.602 n"0.107]10~1 k"0.764]10~5 exp(!22 500/R¹) K "0.110]10~6 CO exp(58 330/R¹) K "0.853]10~2 H2 exp (17 080/R¹) K "0.762]10~1 CO2 exp(26 670/R¹)

Seyfert and Luft (1985) k(/ /2 !/ /K ) CO H2 CH3 OH %2 r" (1#K / #K / #K / #K / #K/ /2 )2 CH3 OH CH3 OH CO2 CO2 CO H2 CO CO H2 H2

A

0.298

k"1.002 exp(!16 670/R¹) K "0.115]10~2 CO exp(18 750/R¹) K "0.475]10~2 exp(3100/R¹) H2 K "0.355]10~2 CO2 exp(12 920/R¹) K"0.131]10~2 exp(17 500/R¹)

p p2 ! CH3 OH (p p1@2)~1.30 (15) CO H2 CO H2 K %2

0.254

k"0.991]10~2 exp(!18 330/R¹)

A

0.176

k"0.830 exp(!23 750/R¹)

Agny and Takoudis (1985) r"k p

p Skrzypek et al. (1985) r"kp1@2p2 1! CH3 OH CO H2 K p p2 %2 CO H2

B

B

(16)

(14)

kp3@2p H2 CO Matulewicz et al. (1984) r" (17) 0.141 (1#K1@2p1@2)(1#K p #K p p1@2#K p p ) H2 H2 CO CO 1 CO H2 2 CO H2

Graaf et al. (1988)

//2 K ) kK (/ /3@2!/ CO CO H2 CH3 OH H2 p (1#K #K )(/1@2#K / /K1@2) CO CO2 H2 H2 O H2 O H2

(18)

dependent on [CO ] in the denominator of the kinetic 2 equation gives rise to a better "t of Kuczynski et al.'s (1987) model. The "t remains, however, poorer than the model of Natta (1955).

0.563]10~2

k"0.112 exp(!147 920R¹) K "0.327]10~9 CO exp(47 500/R¹) K "0.270]10~8 H2 exp(35 000/R¹) K "0.703]10~6 exp(10 830/R¹) 1 K "0.213]10~8 exp(27 920/R¹) 2 k"1346.5 exp(!12 400/R¹) K "0.510]10~2 CO exp(51 670/R¹)

An empirical model was also used for general reference, but did not result in a good "t. This could be explained by omitting the concentration term in [CO ]. However, once 2 the empirical model incorporates [CO ], it becomes 2

J. ErenJ a et al. / Chemical Engineering Science 55 (2000) 1845}1855

Fig. 3. Comparison for t"t between the experimental values of CO %2 conversion (points) and the values calculated with the kinetic model of Schermuly and Luft (1977) (lines), for di!erent values of temperature and space time.

Fig. 4. Comparison for t"t at di!erent temperatures and space %2 times between the experimental values of CO conversion (points) and calculated values using kinetic model of Bakemeier et al. (1970) (lines).

Bakemeier et al.'s (1970) model, which as stated in Table 2 provides an acceptable "t to the experimental data. In conclusion, CO appears to have an important role 2 in the conversion of synthesis gas using this catalyst, due to the fact that CO occupies an important fraction of 2 the methanol synthesis sites. The overall result is an inhibiting e!ect on the rate of synthesis gas conversion. Kinetic models which do not include [CO ] in the deno2 minator of the rate equation are bound to be inadequate. 5. Catalyst deactivation 5.1. Coke deposition Catalyst deactivation is due to the deposition of coke (carbonaceous-like material) on the catalyst surface. The

1851

Fig. 5. Comparison for t"t at di!erent temperatures and space %2 times between the experimental values of CO conversion (points) and the values calculated with the kinetic model of Graaf et al. (1988) (model a) (lines).

coke content was determined by elemental analysis of samples of the deactivated catalyst. Fig. 7 reports the coke content (in wt% referred to the catalyst) as a function of the operating pressure. Results reported correspond to experiments carried out under the following operating conditions: temperature, 4003C; space time, 62.22 (g of catalyst) h (mol(CO#H ))~1; time-on2 stream, 240 min; molar ratio CO/H , 1/2; molar #ow of 2 reactants, 1.15 (mmol of (CO#H )) min~1. 2 As reported in Fig. 7, coke content increases linearly with total pressure until the total pressure reaches 20 atm and then stabilizes at &1.1 wt% of coke at &50 atm. This illustrates the e!ect of the total pressure on the condensation reactions leading to the production of coke. In any case, the coke content remains relatively low relative to the maximum coke capacity of this catalyst. This low coke content is in agreement with the slow rate of catalyst deactivation observed. Furthermore, the elemental analysis indicates that this is a light coke, hardly evolved with an H/C ratio between 1 and 2. A coke of this nature is assigned to a hydrogen-rich atmosphere a!ecting condensation and cyclyzation steps of coke formation. Furthermore, the hydrogenated nature of this coke explains its instability even during the application of the techniques of measurement and analysis. Thus, in Fig. 8, the signi"cant decrease in coke content is reported following a purge with helium (for 90 min) at di!erent temperatures. The sample studied, as shown in Fig. 8, had an initial coke content of 0.95 wt% and once the purge at high temperature was completed, it held only 0.22 wt% of coke. This result was also observed on the coke deposited on HZSM-5 zeolite in the transformation of methanol into gasoline at moderate reaction temperatures (Schulz, Lau & Claeys, 1995; Gayubo, Aguayo,

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Table 3 Proposed kinetic models Kinetic model

Natta (1955) with CO term 2 k(/ /2 2 !/ 3 /K ) CO H CH OH %2 (1#K / #K 2 / 2 #K 3 / 3 #K 2 / 2 )3 CO CO H H CH OH CH OH CO CO

(19)

Error objective function Eq. (1)

Kinetic constants

0.108]10~2

k"0.358 exp(!3400/R¹) K "0.925]10~2 exp(3100/R¹) CO K 2 "0.385]10~1 exp(3000/R¹) H K 2 "0.661]10~1 exp(6200/R¹) CO k"0.112 exp(!11500/R¹)

Matulewicz (1984), modi"ed kp3@2 p H2 CO r" (20) (1#K1@2 p1@2 )(1#K p #K p p1@2#K p p 2 #K 2 p 2 ) CO CO 1 CO H2 2 CO H CO CO H2 H2

Kuczynski et al. (1987), modi"ed k(/ /2 2 !/ 3 /K ) CO H CH OH %2 r" (1#K / #K/ /2 2 #K 2 / 2 ) CO CO CO H CO CO

(21)

Empirical (22) r"kpm p 2 CO H

Fig. 6. Comparison for t"t at di!erent temperatures and space %2 times between the experimental values of CO conversion (points) and the values calculated with the modi"ed kinetic model of Natta (1955) (lines).

Benito, Landeta, Castilla & Bilbao, 1997b) and it demonstrates that modelling catalyst deactivation based on coke-on-catalyst along, may not be provide a rigorous approach. Furthermore, it is rather di$cult to identify those species chemisorbed on the catalyst which also deactivate the catalyst.

0.141

0.521]10~2

0.337]10~1

K "0.361]10~8 exp(8300/R¹) CO K 2 "0.232]10~6 exp(2900/R¹) H K "0.293]10~7 exp(6500/R¹) 1 K "0.208]10~7 exp(4100/R¹) 2 K 2 "12.873 exp(4300/R¹) CO k"0.291 exp(!6000/R¹) K "0.332]10~4 exp(6000/R¹) CO K"0.174]10~7 exp(7900/R¹) K 2 "1.332 exp(8100/R¹) CO k"0.111 exp(!11 100/R¹) m"2.994 n"!0.104

Fig. 7. E!ect of total pressure upon catalyst coke content. Temperature: 4003C; space time: 62.22 (g of catalyst) h (mol (CO#H ))~1; 2 time-on-stream: 240 min; CO/H molar ratio: 0.5; molar #ow of react2 ants: 1.15 (mmol (CO#H )) min~1. 2

Fig. 9 illustrates the e!ect of space time upon coke content. This provides evidence of the greater coke deposition at the inlet of the reactor and of the attenuation of coke deposition towards the reactor outlet, as the concentration of "nal products increases in the reaction medium. These experiments were carried out at 30 atm and under the same aforementioned operating conditions. Fig. 9 also suggests that both primary products or

J. ErenJ a et al. / Chemical Engineering Science 55 (2000) 1845}1855

Fig. 8. E!ect of the sweeping gas temperature upon coke content in the catalyst.

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Fig. 10. Decrease on CO conversion with time-on-stream as a result of catalyst deactivation for di!erent temperatures. Pressure, 30 atm; space time: 62.22 (g of catalyst) h (mol (CO#H ))~1; CO/H molar ratio: 0.5; 2 2 molar #ow of reactants: 1.15 (mmol (CO#H )) min~1. 2

odology for analysing the deactivation measurements consisted of "tting, by non-linear regression, the conversion of CO and its changes with time-on-stream (points in Fig. 10) to the CO conversion resulting from the numerical solution of Eq. (8). On this basis, a kinetic model, non-dependent on the concentration of the components in the product stream, was proposed: da ! "k ad, d dt Fig. 9. E!ect of space time upon coke content in the catalyst. Temperature: 4003C; pressure, 30 atm; time-on-stream: 240 min; CO/H molar 2 ratio: 0.5; molar #ow of reactants: 1.15 mmol (CO#H )) min~1. 2

reaction intermediates are coke precursors. This explanation concurs with recent studies on the deactivation of a HZSM-5 zeolite used in the transformation of methanol into gasoline (Gayubo, Benito, Aguayo, Castilla & Bilbao, 1996; Aguayo, Gayubo, Ortega, Olazar & Bilbao, 1997) where it was proved that coke precursors are mainly oxygenates (methanol and dimethyl ether). Nevertheless, in a reaction such as the one considered in the present study, other explanations may be possible, e.g. close to the reactor's entrance, some Fischer}Tropsch synthesis in parallel with the reaction studied may occur. This may yield organic products with high molecular weight which remain deposited on the catalyst. 5.2. Kinetic modelling of deactivation Experiments were done at di!erent temperatures to quantify the deactivation of the catalyst. The operating conditions were: temperature: 350, 375, 400 and 4253C; space time: 62.22 (g of catalyst) h (mol(CO#H ))~1; 2 time-on-stream: 22 h; CO/H molar ratio: 0.5; molar #ow 2 of reactants: 1.15 (mmol of (CO#H )) min~1. The meth2

(23)

where the variable a, activity, represents the ratio of the following reaction rates: r (24) a" . r 0 The values of the parameters of Eq. (23) which "t the experimental data best are: k "0.0563 exp(!160/R¹); d d"1.0. From these results, it can be appreciated that temperature has a small e!ect, because of the very lowactivation energy. The "t of Eq. (23) to the experimental results is reported in Fig. 10, where the points represent the experimental values. The curves were calculated by solving the mass conservation Eq. (8) in the reactor, with the kinetic parameters which best "t Eq. (23). It is noteworthy that the data in Fig. 10 correspond to times on stream measured after 6 h; this is the time on stream from which deactivation is apparent. Consequently, the data in Fig. 10 have been obtained after those in Fig. 1.

6. Conclusions A catalyst made of a metallic component, Cr O }ZnO, 2 3 and an acidic component, ZSM-5 zeolite, has been

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J. ErenJ a et al. / Chemical Engineering Science 55 (2000) 1845}1855

studied for syngas transformation into liquid hydrocarbons (boiling point in the gasoline range). The measured rates of reaction enable kinetic models to be considered. Kinetic models with term in the partial pressure of CO 2 in the denominator of the rate equation were shown to provide a good "t to the experimental data. The validity of the model proposed by Natta (1955) when there is CO but no methanol in the reaction medium has been 2 proved, for the temperature range 350}4253C. Catalyst deactivation proved to be a slow process compared to the usual deactivation by coke in other processes where ZSM-5 zeolite is employed (e.g. MTG process). This can be explained by the relatively high partial pressures of H inhibiting the development of 2 coke. The rate of catalyst deactivation showed a small e!ect of temperature (i.e. a very low-activation energy) on the deactivation rate and, consequently, the deactivation rate is essentially proportional to the remaining activity of the catalyst.

Acknowledgements This work was carried out with the "nancial support of the UPV/EHU (Project 069.310-EB013/92) and of the University of the Basque Country (Project G34-98).

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