Ind. Eng. Chem. Res. 1991, 30, 2280-2287
2280
Kinetics of Shape-Selective Xylene Isomerization
over
a
ZSM-5 Catalyst
Oscar Cappellazzo,* *Giacomo Cao,1 Giuseppe Messina,1 and Massimo Morbidelli*·1 1
Centro Ricerche Enichem Anic SpA, Zona Industríale Marinella, 07046, Porto Torres (SS), Italy, and Dipartimento di Ingegneria Chimica e Materiali, Universita di Cagliari, Piazza d'Armi, 09123 Cagliari, Italy
The kinetics of xylene isomerization in the liquid phase over a ZSM5 catalyst has been investigated in the temperature range 523-573 K. The presence of shape selectivity toward p-xylene has been examined and classified by comparison with that of other catalysts previously investigated in the literature. Two effective kinetic models, which explicitly account for shape selectivity, have been developed and compared with the obtained experimental data.
Introduction
for liquid-phase isomerization. The obtained data have been reproduced by a kinetic model based on the following triangular reaction scheme:
A classical problem in the petrochemical industry is the treatment of the fraction C8 aromatics for the production of ethylbenzene and xylene isomers. The existing processes are based on two coupled operations: separation (usually through selective adsorption or crystallization for p-xylene and distillation for o-xylene) and isomerization. In the instance where the xylene isomer mixture (free of ethylbenzene) is isomerized, acid catalysts are usually adopted that promote the skeletal rearrangement of the aromatic molecules, operating either in the liquid phase or in the gas phase. If ethylbenzene is also present in the mixture and needs to be isomerized, then it is necessary to operate under hydrogen pressure and to use bifunctional noble metal acid catalysts. In this case the process operates in the gas phase at temperature values ranging between 620 and 720 K. In this work, the previous situation will be considered, with specific reference to liquid-phase isomerization catalyzed by a ZSM5 zeolite at relatively low temperature values (i.e., 520-570 K) in order to minimize catalyst deactivation and side reactions. Silica-alumina catalysts have been widely investigated in the literature, either as such (Hanson and Engel, 1967) or with the addition of 4 wt % nickel (Cortes and Corma, 1978; Corma and Cortes, 1980). Such detailed kinetic studies have shown that the reaction path on silica-alumina catalyst, independently of the presence of nickel which does not affect the isomerization reactions, is the following: ,
ortho
13
^
meta
para
5
ortho
(2)
meta
However, the estimated values for the rate constants of the ortho-para reactions are much smaller than all the others, thus indicating that the prevailing reaction path is the one given by eq 1 also in this case. This conclusion is confirmed by Sreedharan and Bhatia (1987), who represented the gas-phase-isomerization data of m-xylene over a nickel hydrogen mordenite catalyst using reaction scheme 1. In contrast with all the above studies, in the latter one a dual site surface reaction has been adopted as the rate-determining step in the Langmuir-HinshelwoodHoughen-Watson procedure applied for developing the kinetic model. Collins et al. (1982) studied the kinetics of xylene isomerization over a LaY zeolite and concluded that the mutual interconversion between o- and p-xylene is quite difficult, thus again indicating eq 1 as the prevailing reaction path. Considerable improvement in the p-xylene selectivity of the isomerization process has been obtained with the use of small pore zeolites such as ZSM5. Collins et al. (1983) and Young et al. (1982) have used HZSM5 for xylene isomerization in the liquid phase and in the gas phase, respectively. In both cases a shape-selective effect toward p-xylene has been evidenced by showing that pxylene is produced immediately at very low conversion values from m-xylene as well as from o-xylene. This behavior, which is not accounted for by kinetic scheme 1, is distinctly different from the one discussed above for silica-alumina or large pore zeolite catalysts (e.g., hydrogen mordenite or LaY). An even stronger shape selectivity toward p-xylene is exhibited by suitably modified ZSM5 catalysts, whose channel dimensions are further reduced through the introduction of compounds such as phosphorus and magnesium (Young et al., 1982). The proposed explanation for this behavior is in the occurrence of diffusion limitations in the transport of the xylene isomer molecules in the ZSM5 channels, whose dimension (ca. 6 Á) is comparable to the molecular dimension (Young et al., 1982). Olson et al. (1981) have determined that the diffusion coefficients in ZSM5 for oand m-xylene are about 103 times lower than that for pxylene. Thus, p-xylene would be able to immediately leave the zeolite channels, as soon as formed, while o- and mxylene are retained for longer times thus leading to the possibility of further isomerization. In other words, the reaction inside the zeolite channels occurs again via intramolecular 1,2 shifts of the methyl groups, thus in ac-
(1)
which proceeds via intramolecular 1,2 shifts of the methyl groups, thus allowing the transformation of o- into p-xylene (or vice versa) only through m-xylene as an intermediate step and not directly as a single step (Cortes and Corma, 1978). Accordingly, a rather good interpretation of the experimental isomerization data has been obtained by Corma and Cortes (1980) using a kinetic model developed through the Langmuir-Hinshelwood-Houghen-Watson procedure assuming as the rate-controlling step a single site surface reaction. This kinetic model has been successfully applied by Bhatia et al. (1989) for simulating the behavior of an adiabatic fixed bed catalytic reactor. The silica-alumina catalyst used in the latter work contained 0.5 wt % platinum in order to convert the ethylbenzene present in the feed into the xylene isomers (cf. Hsu et al., 1988).
Also zeolite catalysts have been considered for xylene isomerization. Hydrogen mordenite has been used by Hopper and Shigemura (1973) and Norman et al. (1976) To whom correspondence should be addressed. Centro Ricerche Enichem Anic SpA. * Universita di Cagliari. * 1
0888-5885/91/2630-2280$02.50/0
para
©
1991 American Chemical Society
Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991
2281
COOLER
-(
>·
A n JJ
LIQUID
PRODUCT
LIQUID XYLENE
L_ w
HEATER
Figure
1.
Schematic diagram of the experimental apparatus.
cordance with kinetic scheme 1. However, the presence of diffusional limitations much more pronounced for o- and m-xylene than for p-xylene leads to a multistep interconversion in the zeolite channels, from which the smaller p-xylene has a much higher probability of diffusing in the bulk phase. The net result of this process is the apparent formation of p-xylene directly from o-xylene, thus in agreement with reaction scheme 2. A quantitative mathematical model reproducing the above mechanism has been developed by Wei (1982), who has shown that indeed the occurrence of diffusional limitations shifts the composition trajectories in the classical triangular plot from those typical of non-shape-selective catalysts to those typical of the shape-selective ones. More detailed studies, based on a Monte Carlo model, have confirmed that the presence of pore-blocking elements (such as phosphorus or magnesium in ZSM5) can enhance the para shape selectivity of the catalyst (Theodorou and Wei, 1983; Sundaresan and Hall, 1986). Turner and Ramkrishna (1986) have considered Wei’s model in the context of a pseudohomogeneous axial dispersion reactor model. In the case of first-order kinetics the problem is linear and can be solved analytically through linear operator theory. The calculated results indicate that the shape-selective character of the catalyst leads to a strong dependence of the reactor outlet selectivity on feed composition, residence time, and axial dispersion. The occurrence of shape selectivity toward p-xylene formation has been further investigated in the literature in more recent years. Chen and Gardwood (1986) have extensively reviewed the industrial application instances of shape-selective catalysts. The possibility of using this phenomenon as a tool for catalyst characterization has been examined (cf. Dewing, 1984; Martens et al., 1988). Ratnasamy and Pokhriyal (1989) have found that shape selectivity is enhanced by silylating the active sites on the external surface of the zeolite crystals. This finding indirectly supports the mechanism reported above, since it can be explained by considering that such sites, being located on the zeolite crystal surface, cannot participate in the selective diffusion-reaction process. The aim of this work is to study the isomerization reactions in liquid xylene isomer mixtures on a commercial Mobil ZSM5 isomerization catalyst, which exhibits some shape selectivity toward p-xylene formation. The main characteristics of such a selectivity are classified by comparison with the composition trajectories in the triangular plot arising when using other catalysts previously inves-
tigated in the literature. Moreover, two kinetic models are proposed and compared with the experimental data. The first one is based on kinetic scheme 2, which is regarded as an effective tool for simulating the para-selective behavior of the reacting system. In the second one, along the lines of Wei’s 1982 work, kinetic scheme 1 is adopted and the reaction-diffusion process in the zeolite pores is simulated through a lumped mass transport model in order to account for the shape selectivity phenomenon. Both models have been kept as simple as possible from the mathematical point of view so as to make possible their application in reactor design and optimization.
Experimental Apparatus and Procedure A sketch of the experimental apparatus used in this study is shown in Figure 1. The jacketed tubular reactor, constituted by a fixed bed (0.031-m diameter and 0.33-m length) containing 105 g of catalyst pellets (0.0016 m X 0.004 m), is maintained at the desired temperature through forced oil circulation. The reactant feed stream is fed to the reactor, at the desired composition, through a dosimetric pump and a preheating section. The system is also equipped with a back-pressure regulator in the effluent line which is connected to the nitrogen line. Before starting the reaction, the system is pressurized up to 30 atm in order to keep the reacting mixture in the liquid phase at the desired reaction temperature. During each experimental run, the reactor effluent is first brought to the ambient temperature in a water cooled jacket tank, maintained at the working pressure, and then drained into an atmospheric pressure container through a pneumatic valve. The outlet stream is analyzed via a gas chromatograph (Carlo Erba 4200), equipped with a 2-m steel capillary column (0.00315-m diameter) filled with Chromosorb 80/100. Catalyst activity is tested every 50 h of operation by repeating a given experimental run at standard operating conditions: no significant deactivation has been found during the entire experimental investigation. During each experimental run, reactor steady-state conditions and catalytic bed isothermality have been carefully checked by moving a thermocouple along the bed axis. Various values of the reactor residence time have been investigated by changing the feed flow rate with the dosimetric pump mentioned above.
Kinetic Models Before presenting the kinetic models and the experimental runs performed to estimate the corresponding ad-
2282
Ind. Eng. Chem. Res., Vol. 30, No.
10, 1991
Table I. Data and Relationships Used for the Estimation of Transport Limitations (First Experiment in Table II, Run 1)
« fb
0.033 cm 0.93 g/cm3 0.90 g/cm3 0.0085 mol/cm3 0.4 0.4
k, T
6.1 cm/h 553 K
rexp
0.072
d¡
3.1 cm
Q
2333 cm3/h 20 h'1 0.058 cm2/h
Vp/sp #>P
Pi
^eip
LWHSV De
mol/[h g(catalyst)]
[30 coth (30)
V
-
l]/302
justable parameters, it is necessary to worry about the presence of transport resistances that may mask the experimental reaction rate data. While deviations from isothermality conditions can certainly be neglected due to the modest heat involved in the isomerization reactions under examination and the efficient heat transfer guaranteed by the cooling jacket, the effect of mass-transfer limitations requires a careful examination. In Table I are summarized the physicochemical parameter values and the relationships used for evaluating the mass-transfer coefficients. In particular, the experimental run corresponding to the highest value of the reaction rate has been considered, in order to obtain an estimate of the maximum concentration gradient that may arise in the entire experimental analysis. The maximum difference between the reactant concentrations at the catalyst surface and in the bulk stream can be estimated as follows (cf. Carberry, 1976):
explain the para shape selectivity exhibited by the catalyst under examination, as has been shown by the models proposed by Wei (1982) and Theodorou and Wei (1983). Such models are indeed very useful for explaining the mechanistic aspect of the process, but they are not practical for reactor design purposes. In industrial reactors inter- and intraphase mass and heat transports are usually important. Thus, a reliable model of such units should include these aspects and in particular the diffusion-reaction process in the catalyst particle macropores. When this is done by using a detailed kinetic model, such as those mentioned above, a reactor model constituted by a system of three-dimensional partial differential equations is obtained, which involves three characteristic lengths of the system: reactor length, particle diameter, and zeolite pore length. In an attempt to simplify the reactor model, we explore the possibility of deriving simplified effective kinetic models that can describe the shape-selective behavior of the catalyst. One possible way of doing this is to consider the triangular kinetic scheme (2), which has the possibility over scheme 1 of effectively accounting for the para-shape-selective behavior of the catalyst (e.g., the production of p-xylene from pure o-xylene immediately at very low conversion values). A second possibility, based on a lumped description of the diffusion-reaction process in the zeolite channels, is examined next. Assuming a single-site, surface rate controlled, Langmuir-Hinshelwood kinetics, the following expressions for the rates r, of each one of the interconversion reactions in scheme 2 and for the corresponding kinetic constants, k¡ are obtained: can
=
fcxCo/D;
k\
=
k\K0aa
(5)
r2
=
k2CM/D;
k2
=
k'2KMa0
(6)
(3)
r3
=
k3Cw/D\
k3
=
k'3KMa0
(7)
where rexp represents the measured value of the reaction rate, Vp and Sp are the catalyst particle volume and external surface, and k( is the interparticle mass transfer coefficient estimated through appropriate semiempirical relationships. Since the obtained value of the difference AC is equal to about 4% of the reactant concentration in the feed stream (cf. C„p in Table I), it can be concluded that interphase transport limitations can be neglected in the analysis of the experimental data. On the other hand, the presence of concentration gradients within the catalyst particle can be estimated through the following relationship (cf. Carberry, 1976):
r4
k4Cp/D\
k4
=
k'4KpC0
(8)
&&
(9)
AC
=
rexppp(Vp/Sp)/fef
2
=
=
0.00036 mol/cm3
rMpPp(Vp/Sp)2/DeCexp
=
0.15
(4)
is the normalized Thiele modulus, which is related to the effectiveness factor through the approximate relationship reported in Table I (cf. Aris, 1965). Substituting such relationship in eq 4 leads to = 0.41, and then = 0.91. This shows that in the most drastic conditions considered in the experimental analyis, intraphase concentration gradients start to be significant. However, considering that this situation occurs only for a very few experimental data, at the lowest conversion values, also intraphase concentration gradients have been neglected in the following analysis, thus assuming that all the experiments have been performed in the regime controlled by the chemical reactions. As mentioned in the Introduction, the mechanism of xylene isomerization in small pore zeolites involves simultaneous diffusion and reaction in the zeolite channels. The coupling of kinetic scheme 1 with the diffusion process, which strongly favors p-xylene over the other isomers,
where
=
rb~ re
=
~
^6 Cp/D;
k3
(10)
k'gKpC0
—
where D 1 + KpCp + KMCM + K0C0\ the indices O, M, and P refer to the three isomers o-, m-, and p-xylene, respectively; the indices 1-6 refer to the reactions shown + CP* + CM* + C0* represents the in eq 2; and 0 = overall concentration of active sites in the catalyst. Following the conclusion drawn above about the role of =
transport limitations, the behavior of the experimental fixed bed catalytic reactor has been simulated through an isothermal one-dimensional, pseudohomogeneous, plugflow model. Thus, the mass balances for m- and p-xylene are given by (11) dCM/dr = r1 + r4-r3-r2
dCp/dr
r3 + r5
=
-
r6
-
(12)
r4
with initial conditions C, C,° at 0, while the o-xylene concentration is readily obtained through the algebraic overall mass balance =
C0
=
r
Cm0 + C0° + CP°
=
-
Cm
-
CP
(13)
It is worth noticing that catalyst deactivation as well as the production of byproducts, such as ethylbenzene or toluene, are not accounted for in the model in accordance with experimental findings. An important aspect in developing a reliable kinetic model for the reacting system under examination is its thermodynamic consistency at equilibrium. This condition can be introduced in the model by enforcing that at equilibrium conditions the net rate of each reversible re-
Ind. Eng. Chem. Res., Vol. 30, No. 10,1991 action in scheme 2 vanishes, thus obtaining the following relationships: k2 =
ki/KM0
(14)
k3KM0/KPQ
(15)
=
=
Table II. Operating Conditions for the Performed Experimental Runs_ feed composn, wt % run temp, K 2 3
4
and KM0 = (CM/C0)eq are the (CP/C0)eq the o- to p-xylene and o- to constants for equilibrium m-xylene reactions, respectively. The values of such constants can be readily computed from the thermodynamic data of free energy of formation reported by Stull et al. where KP0
=
5
6 7
8 9 10
(1969).
The influence of temperature on the model kinetic parameters is accounted for through the following Arrhenius-type relationships: Ki = K¡(T0) exphfí.U/ 1/T0)/R], i =0, , P -
(17)
kj
=
kj(T0)
exp[-£y(l/T l/T0)/fi], j -
=
1-6
(18)
where T0 is a reference temperature within the range of values investigated experimentally. Thus, considering constraints 14-16, the model contains three rate constants and three adsorption constants, which, accounting for their temperature dependence, lead overall to 12 adjustable parameters that have to be estimated by comparison with the experimental data. As an alternative approach to the simulation of the shape-selective isomerization, a kinetic model has been developed that explicitly accounts for the transport process in the zeolite channels coupled with the isomerization reactions. In order to keep the model as simple as possible, the transport process has been simulated by a lumped mechanism, through the introduction of suitable masstransport coefficients (Glueckauf, 1955). Moreover, it is assumed that the three isomers inside the zeolite channels, whose concentration values are denoted by the subscript c, undergo the isomerization reactions via intramolecular 1,2 shifts of the methyl groups, thus following reaction scheme 1. Accordingly, the kinetic model considers the
following scheme: *
(ortho),
(meta),
(para).
(19)
4
2
|
t
t
ortho
meta
para
1
1
where the vertical arrows represent the mass-transport process. Assuming pseudo-steady-str*e conditions for the isomer concentrations in the zeolite channels the following mass balances are obtained: (20) 0m(Cm CMc) = (r2 + r3-r1- r4)pP '
dP(CP
-
CPc)
=
(r4
r3)pP
-
(21)
(22) 0o(Cq Cq,) (fi r2)pP where r; is the rate of the jth reaction, as defined by eqs 5-8 for j = 1-4, which is now evaluated using the concentration values in the zeolite channels (i.e., Cic), while /?,· represents the ratio between the mass-transport coefficient of the ith component and the characteristic length of the transport process. The evaluation of these quantities “a priori” is not feasible, so that a fitting procedure of the available experimental data is needed. For the reaction rates, this involves the evaluation of 2 kinetic constants (since the equilibrium constraints 14 and 15 are used for the other two kinetic constants) and 3 adsorption constants, which, using eqs 17 and 18 in order to account for the temperature dependence, lead overall to 10 adjustable parameters. -
—
-
553 553 553 553 523 523 523 573 573 573
1
(16)
^6-^PO
2283
P
=
O
=
0
=
P
=
O
=
P O
=
=
=
=
=
100% 100% 100% 27.5%, M 100% 100% 100% 100% 100% 100%
About the mass-transport parameters, d¿, their values proportional to the corresponding diffusion coefficients in the zeolite channels (Glueckauf, 1955). Thus, in order to keep low the number of adjustable parameters, we take advantage of the measured values of such coefficients reported by Olson et al. (1981) and assume accordingly that are
ß
=
ß0
=
10~3ß
(23)
Moreover, an Arrhenius-type dependence of ß from temperature has been introduced (cf. eq 18). This leads overall to 12 adjustable parameters as in the previous case of the triangular kinetic model. It should be mentioned that the value of the adjustable coefficient ß cannot be predicted by using the usual empirical relationships for estimating transport coefficients, since it refers to a process occurring in the zeolite channels (at a molecular level), and therefore it has to be regarded as a parameter characteristic of the catalyst that determines its intrinsic kinetics. The reactor model is again given by eqs 11-13 above, with r5 = r6 = 0, recalling that the reaction rates have to be evaluated at the concentration values, Cic. These are obtained from the bulk values C¡, by solving at each time instant the system of eqs 20-22.
Comparison with Experimental Data The operating conditions at which the experimental runs have been performed are summarized in Table II. Three temperature values (523, 553, and 573 K) and various feed compositions are investigated at several values of the liquid
weight hourly space velocity (LWHSV), ranging from 1.85 to 20 g(feed)/[h g(catalyst)]. Because most of the experimental data are available at 553 K, this has been chosen as the reference temperature (i.e., T0 = 553 K in eqs 17 and 18). Because of the limited number of experimental data available, it appeared convenient to reduce the number of adjustable parameters by eliminating the adsorption parameters KM, KP, and KQ, thus setting D = 1 in eqs 5-10. In the case of the second model above, the system of eqs 20-22 can be solved analytically leading to the following explicit expressions for the corresponding reaction rates:
r\
=
felCo,
(24)
r2
=
k2CMc
(25)
r3
=
fe3CM,
(26)
r4
=
fe4CP,
(27)
where
c
_
d'pCP + M'mCm»'1 + k4 +
ß'
klk3^0C0al{kx + do)'1 -
k3k±arl (28)
£ Mc
_
0'mGm + k2 +
M'oCo(fei + do)-1 + k,CPc
k3-
fe1A2(fe1
+ d'o)"1 + 0'm
2284
Ind. Eng. Chem. Res., Vol. 30, No. Cqc
with a i = ,
=
k2 + k3
O,
.
-
=
ß'
10, 1991
+
Ai +
(30)
ß'
kxk2{kx + ß' )'1 +
Table III. Parameter Values Obtained through the Fitting Procedure for Model 1 (Eqs 5-13, 0=1) and Model 2 (Eqs 11-13 and 24-30)
ß'
and
/?',·
=
/3,·/ ,
For both models mentioned above, a nonlinear regression procedure, using the sum of the squared deviations between calculated and measured weight fractions of all the involved species as objective function, has been used to estimate the model parameters and fit the experimental data. In order to expedite the fitting procedure, first only the data at 553 K have been considered (i.e., runs 1-4 in Table II), thus estimating the three adjustable model parameters in the isothermal model simulation at T = 553 K [i.e., kj(T0) with j = 1, 3, 6 and /3P], The obtained parameter values are summarized in Table III for each model together with the corresponding values of the average percentage error. Next, the experimental data at 523 and 573 K have been considered (runs 5-10 in Table II) to estimate the values of the remaining parameters accounting for temperature effects (i.e., the activation energies Ej with j = 1, 3, 6 and Eg). Again the obtained results are summarized in Table
parameters
model
*i(T0), cm3/h
2.29 3.80 2.08
g
k3(T0), cm3/h g k6(T0), cm3/h g j8p. h"1
e%
Ex, Es, Eg, Ee,
(%
kcal/mol kcal/mol kcal/mol kcal/mol
6.1 23.6
32.5 37.8 8.2
1
model
2
0.122 0.0844
7957.0 6.1 32.9
36.6 28.0 8.3
O^thO
III.
As mentioned above, the two adopted models are, even though to a different extent, simplified effective models, which do not attempt a detailed description of the involved elementary processes. As a consequence, it is not surprising that the parameters kx, k2, Ex, and E2 in Table III take different values in the two models, even though in principle they refer to the same physical process. For this same reason, a direct comparison of the values of such parameters with those reported in previous studies is not possible, particularly when considering that they all used substantially different catalysts. However, it can be observed that the values of the activation energies are similar to those reported by Ma and Savage (1987) for H-USY ultrastable faujasite, while Corma and Cortes (1980) and Hanson and Engel (1967) reported for a silica-alumina catalyst values that are substantially lower than those obtained in this work. A direct comparison between the predictions of the first kinetic model (i.e., eqs 5-13) and the experimental data in terms of weight percentages of the involved species in the reactor outlet is reported in Table IV. In Figure 2 are shown in the classical triangular plot the trajectories followed by the composition of the reacting mixture: the points represent the experimental data while the continuous curves represent the calculated values. The good agreement observed indicates that the proposed kinetic model is suitable for design purposes. A similar conclusion can be drawn for the second kinetic model (i.e., eqs 24-30), whose accuracy in reproducing the experimental data is very similar to that of the first model, as it appears from the values of the percentage errors reported in Table III. As a point of comparison, the same experimental data have also been simulated by using a kinetic model based purely on reaction scheme 1. Such a model constitutes the limiting behavior of the first model for ks = k6 = 0, as well as of the second model for /3P = °°. It includes overall only four adjustable parameters (using equilibrium constraints 14 and 15). As expected, the agreement with the experimental data significantly deteriorates with respect to the previous cases, leading to values of the average percentage error equal to 17.2% for runs 1-4 in Table II and 19.8% for the others. It is interesting to note that this model exhibits the largest error (about 80%, as opposed to the triangular model whose errors never exceed 30%) in the experimental run relative to the lowest residence time and
Figure 2. Comparison between computed trajectories (continuous curve) and experimental data (O) for runs 1-4 in Table II; *, equilibrium composition at 553 K. Ortho
ture composition obtained using different catalysts: (O) PZSM5 (T = 873 K), Young et al., 1982; ( ) MgZSMS (T = 823 K), Young et al., 1982; ( ) HZSM5 (T = 573 K), Young et al., 1982; (V) HZSM5 (T = 523 K), Collins et al., 1983; (0) ZSM5 (T = 553 K), this work; (·) SiOjAljOa (T = 673 K), Cortes and Corma, 1978; ( ) LaY (T = 623 K), Collins et al., 1982; ( ) mordenite (T = 505 K), Hopper and Shigemura, 1973;(©)equilibrium (T = 523 K).
feed constituted by pure p-xylene at 553 K.
Discussion and Concluding Remarks In order to establish quantitatively the characteristic of the examined catalyst with respect to para shape selectivity, it is convenient to compare the trajectories of the reacting mixture composition in the classical triangular plot obtained with different catalysts. Such a comparison is
Ind. Eng. Chem. Res., Vol. 30, No. 10,1991
2285
Table IV. Comparison between Calculated (Wc) and Experimental (W,) Weight Percentage of m- and p-Xylene in the Reactor Outlet Stream. Operating Conditions as in Table II run LWHSV WPc WMc Wp. Wm. 1 1 1 1 1
2 2 2 2
2 3
3 3
3 3
4 4 4
4 4 5 5 5 6 6
20.0 10.0 6.7 2.5 1.7 20.0 10.0 4.7 2.5 1.8
20.0 10.0 4.7 2.5 1.7 20.0 10.0 4.7 2.5 1.7 5.0 1.7
0.92 3.0 1.0
6
0.55
7 7
1.0 0.55 16.7 8.3 5.0 10.5 5.5 10.0 0.55
8 8
8 9
9 10 10
32.8 43.9 50.0 52.9 52.6 15.9 29.2 42.0 49.5 51.2 84.3 74.6 61.9 55.6 54.8 56.6 54.6 52.8 52.8 52.3 28.9 45.3 48.1 18.8 33.1 42.5 70.8 64.1 46.5 53.8 54.1 41.2 48.1 62.6 55.2
28.0 42.2 49.2 54.7 54.2 10.9 19.5 33.5 46.0 50.0 83.0 73.2 62.5 56.5 54.7 56.5 54.7 53.5 53.5 53.7 22.4 44.3 52.9 11.7 27.1 38.7 69.5 61.1 50.2 54.3 54.0 25.2 39.0 66.4 59.3
60.0 44.4 34.9 25.1 24.6 7.40 13.8 19.6 22.4 23.2 11.9 18.0 23.3 24.7 24.4 16.9 19.9 22.8 23.3 23.4 65.7 43.3 37.2 6.85 12.5 17.2 18.7 21.2 39.8 26.8 24.1 18.4 21.3 22.4 23.5
64.4 45.6 35.6 24.6 24.0 8.52 13.2 18.7 22.0 23.0 12.4 18.6 23.4 24.2 24.0 17.5 20.4 22.9 23.7 23.8 72.7 44.4 31.3 5.80 12.6 17.5 19.4 23.2 32.5 24.9 23.9 17.9 21.4 22.3 24.0
shown in Figure 3, where the experimental trajectories obtained by various authors, using different catalysts and operating either in the liquid phase or in the vapor phase at different temperature values, are reported. The asterisk in Figure 3 represents the equilibrium composition evaluated at T = 553 K, while the closed symbols refer to the non-shape-selective catalysts (such as silica-alumina, mordenite, and LaY), and the open symbols refer to catalysts that exhibit various degrees of shape selectivity. The latter include HZSM-5 and the catalyst used in this work, together with the most selective ones constituted by ZSM-5 catalysts containing phosphorus and magnesium. The extent of shape selectivity for each catalyst can be readily judged from the extent of deviation of the corresponding reaction trajectories from those obtained using the classical non-shape-selective catalysts. The two models developed above provide a substantially equivalent fitting of the obtained experimental data, and both can be used for the optimal design of a large-scale reactor. However, the difference in the way such models account for the shape selectivity is substantial. The first one adopts the triangular kinetic scheme (2), thus adding to the actual mechanism (1) a fictitious reaction, i.e., the direct o-xylene-p-xylene isomerization, in order to account for the p-xylene shape selectivity. The second one, following the mechanism proposed by Wei (1982), considers the actual reaction scheme (1), and includes a masstransport resistance for reproducing the effect of shape
selectivity. The parametric behavior of the two models is illustrated in Figure 4a and Figure 4b, respectively, where it appears that both of them can simulate the deviations in the trajectories of the reacting mixture composition, induced by
4. Effect of shape selectivity on the trajectories of the re= 0.0 acting mixture composition as predicted by model 1: (---) cm3/ [h g(catalyst)], other parameters as in Table III; (—) parameters as in Table III; (—) ft6 = 3.33 cm3/[h g(catalyst)], other parameters as in Table III. (b) Effect of shape selectivity on the trajectories of the reacting mixture composition as predicted by model 2: (- -) dp = and dp· A similar improvement is also obtained by in-
Ind. Eng. Chem. Res., Vol. 30, No.
2286
10, 1991
as adjustable parameters the adsorption constants KM, KP, and K0 in both models. However, at least for the set of experimental data considered in this work, the obtained reduction in the average error is modest, while the correlation among the various adjustable parameters becomes very strong so as to make not reliable the correspondent estimated values. Finally, it is worthwhile stressing the role that the shape-selective behavior of the catalyst may have in the optimal design of a fixed-bed reactor. In particular, since the kinetics of formation of p-xylene in the isomerization process is strongly enhanced, it is quite possible that, starting from pure m-xylene, the reacting mixture composition trajectory goes through concentration values of p-xylene that are higher than the final one correspondent to the equilibrium conditions. This poses a problem of optimal design of the fixed-bed reactor whose length should be selected so as to obtain such a maximum concentration, thus stopping the reaction before equilibrium conditions are reached in order to maximize the process profit. As an illustrative example, using the triangular kinetic model eqs 5-10 and the reactor model eqs 11-13, the problem of determining the optimal values of residence time and operating temperature that maximize the pxylene excess over the equilibrium value, with feed constituted by pure m-xylene, has been solved. It has been found that for = 0.40 h g(catalyst)/cm3(feed) and T = 548 K the p-xylene concentration is larger than the final equilibrium value by about 2%. This result is obviously limited to the particular catalyst considered in this work. Higher values would be obtained by using catalysts exhibiting stronger shape selectivity towards p-xylene.
troducing
Acknowledgment We acknowledge the Mobil Oil Corporation for providing the catalyst used in this work. Part of this work was financially supported by the National Research Council (CNR): Progetto Finalizzato Chimica Fine.
Nomenclature reactant concentration value used in Table I, mol/cm3 concentration of the ith isomer, mol/cm3 = concentration of the ith isomer in the zeolite channels, C¡c mol/cm3 C¡* = concentration of the adsorbed ith isomer, mol/cm3 dt = tube diameter, cm = effective diffusivity, cm2/s De Ei = activation energy of the jth reaction, kcal/mol Hi = heat of adsorption of the ith isomer, kcal/mol kf = interparticle mass transfer coefficient, cm/s kj = kinetic constant of the jth reaction, cm3/ [h g(catalyst)] kj = surface kinetic constant of the ;th reaction, cm3/[h g(catalyst)] K¡ = adsorption equilibrium constant of the ith isomer, cm3/mol KP0 = equilibrium constant of the isomerization reaction of o-xylene to p-xylene KMo = equilibrium constant of the isomerization reaction of o-xylene to m-xylene LWHSV = Qpt/P, liquid weight hourly space velocity, g (feed)/[h g (catalyst)] P = reactor catalyst load, g Q = volumetric flow rate, cm3/h = reaction rate of the jth reaction, mol/[h g(catalyst)] r¡ = measured value of the reaction rate, mol/[h g(catalyst)] reip T = temperature, K T0 = reference temperature, K S„ = catalyst particle external surface, cm2 = catalyst particle volume, cm3 Vp = weight percentage of the ith isomer Vv, Ceip = C¡
=
Greek Letters k2 +
=
a
Pi
k3- kik^kx
+ /3'q)"1
=
P/PP ratio between the mass-transport coefficient of the ith component and the characteristic length of the transport process, h"1 e% = average relative percentage error = bed void fraction eb -
=
ep = =
p¡
=
Pp =
=
0 r
=
particle porosity
effectiveness factor average fluid density, g/cm3 catalyst particle density, g/cm3 concentration of catalyst free sites, mol/cm3 overall concentration of catalyst active sites, mol/cm3 pf/LWHSV, contact time, h g(catalyst)/cm3(feed)
Subscripts c = zeolite channel eq = equilibrium conditions M = m-xylene O = o-xyiene P = p-xylene
Superscript ° inlet conditions =
Registry No. o-Xylene, 95-47-6; m-xylene, 108-38-3; p-xylene, 106-42-3.
Literature Cited Aris, R. Normalization for the Thiele Modulus. Ind. Eng. Chem. Fundam. 1965, 4, 227. Bhatia, S.; Chandra, S.; Das, T. Simulation of the Xylene Isomerization Catalytic Reactor. Ind. Eng. Chem. Res. 1989, 28, 1185. Carberry, J. J. Chemical and Catalytic Reaction Engineering·, McGraw-Hill: New York, 1976; pp 205-224. Chen, N. Y.; Gardwood, W. E. Industrial Application of Shape-Selective Catalysis. Catal. Rev. Sci. Eng. 1986, 28, 185. Collins, D. J.; Mulrooney, K. J.; Medina, R. J. Xylene Isomerization and Disproportionation over Lanthanum Y Catalyst. J. Catal. 1982, 75, 291.
Collins, D. J.; Medina, R. J.; Davis, B. H. Xylene Isomerization by ZSM-5 Zeolite Catalyst. Can. J. Chem. Eng. 1983, 61, 29. Corma, A.; Cortes, A. Kinetics of the Gas-Phase Catalytic Isomerization of Xylenes. Ind. Eng. Chem. Process Des. Dev. 1980,19, 263.
Cortes, A.; Corma, A. The Mechanism of Catalytic Isomerization of Xylenes: Kinetic and Isotopic Studies. J. Catal. 1978, 51, 338. Dewing, J. Are Shape-Selective Reactions Tools for Zeolite Charac-
terisation? J. Mol. Catal. 1984, 27, 25. Glueckauf, E. Theory of Chromatography. Part X—Formulae for Diffusion into Spheres and their Applications to Chromatography. Trans. Faraday Soc. 1955, 51, 1540. Hanson, K. L.; Engel, A. J. Kinetics of Xylene Isomerization over Silica-Alumina Catalyst. AIChE J. 1967,13, 260. Hopper, J. R.; Shigemura, D. S. Kinetics of Liquid Phase Xylene Isomerization over H-Mordenite. AIChE J. 1973,19, 1025. Hsu, Y. S.; Lee, T. Y.; Hu, H. C. Isomerization of Ethylbenzene and m-Xylene on Zeolites. Ind. Eng. Chem. Res. 1988, 27, 942. Ma, Y. H.; Savage, L. A. Xylene Isomerization Using Zeolites in a Gradientless Reactor System. AIChE J. 1987, 33, 1233. Martens, J. A.; Perez-Pariente, J.; Sastre, E.; Corma, A.; Jacobs, P. A. Isomerization and Disproportionation of m-Xylene Select!vities Induced by the Void Structure of the Zeolite Framework. Appl. Catal. 1988, 45, 85. Norman, G. H.; Shigemura, D. S.; Hopper, J. R. Isomerization of Xylene over Hydrogen Mordenite. A Comprehensive Model. Ind. Eng. Chem. Prod. Res. Dev. 1976, 15, 41. Olson, D. H.; Kokotailo, G. T.; Lawton, S. L.; Meier, W. M. Crystal Structure and Structure-Related Properties of ZSM-5. J. Phys. Chem. 1981, 85, 2238. Ratnasamy, P.; Pokhriyal, S. K. Surface Passivation and Shape Selectivity in Xylene Isomerization over ZSM-48. Appl. Catal. 1989, 55, 265.
Sreedharan, V.; Bhatia, S. Vapour Phase Isomerization Study of m-Xylene over a Nickel Hydrogen Mordenite Catalyst. Chem. Eng. J. 1987, 36, 101.
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fusion in a Tubular Reactor: An Operator-Theoretic Solution. Ind. Eng. Chem. Fundam. 1986, 25, 258. Wei, J. A Mathematical Model of Enhanced para-Xyiene Selectivity in Molecular Sieve Catalysts. J. Catal. 1982, 76, 433. Young, L. B.; Butter, S. A.; Kaeding, W. W. Shape Selective Reactions with Zeolite Catalysts. J. Catal. 1982, 76, 418.
Stull, D. R.; Westrum, E. F.; Sinke, G. C. The Chemical Thermodynamics of Organic Compounds·, Wiley: New York, 1969; pp 368-369. Sundaresan, S.; Hall, C. K. Mathematical Modelling of Diffusion and Reaction in Blocked Zeolite Catalysts. Chem. Eng. Sci. 1986,41, 1631.
Theodorou, D.; Wei, J. Diffusion and Reaction in Blocked and High Occupancy Zeolite Catalysts. J. Catal. 1983, 83, 205. Turner, B. G.; Ramkrishna, D. Multicomponent Reaction and Dif-
Received for review October 10,1990 Accepted June 14, 1991
Partial Oxidation of Methane to Methanol in Elevated Pressure
a
Flow Reactor at
Daniel W. Rytz and Alfons Baiker* Department of Chemical Engineering and Industrial Chemistry, Swiss Federal Institute of Technology, ETH-Zentrum, CH-8092 Zurich, Switzerland
The partial oxidation of methane was carried out in a glass-lined tubular flow reactor packed with quartz chips. Experiments were performed at pressures of 20-50 bar, temperatures of 425-500 °C, space times of 3-12 s, and oxygen concentrations in the feed of 3,5, and 10 mol %. Reaction products observed were methanol, carbon monoxide, and water with small amounts of carbon dioxide, ethylene, and formaldehyde. Depending on the experimental conditions, selectivities to methanol of up to 50% could be reached with methane conversions of 3-4%. At higher conversion levels of 14% the selectivity to methanol dropped to 30%. Low temperature and high pressure were found to favor methanol production. Increase of the oxygen concentration in the reactant feed increases the yield but lowers the selectivity to methanol.
Introduction
(1985), and Pitchai and
The world consumption of methanol today is in excess of 25 million tons per year and growing steadily. This indicates that methanol is still gaining importance in the bulk chemical market. More than one-third of the pro-
view
duced methanol is used to make formaldehyde; the rest is mainly used for the production of acetic acid and gasoline octane improvers such as methyl-ieri-butyl ether. Promising new areas for methanol are the direct use in fuels or in fuel cells as well as feedstock for Mobil’s MTG (methanol-to-gasoline) process. The existing commercial process for methanol production involves two steps. In the first step hydrocarbons are re-formed to synthesis gas at high temperatures. This step is highly endothermic, so that some of the hydrocarbons have to be burned in order to run the process. The second step is the low-pressure methanol synthesis, where the syngas is catalytically converted to methanol at around 250 °C and 50 bar. The hydrocarbon sources for the first step include natural gas, petroleum residues, naphtha, and coal, the first being by far the most important. However, in spite of this, natural gas is still a large, underutilized energy resource, mainly because the known reservoirs are located in remote areas (e.g., Siberia, western Canada, offshore Australia) from where the cost of transportation to industrial areas still prohibits the use of the resource. The aim of this work in the area of direct partial oxidation goes in two directions. The first is to minimize the engineering complexity and maximize the energetic efficiency. The second is to develop a process that is simple enough to install at faraway wellheads so that the resources of natural gas can be used and transported efficiently. From the above, it becomes evident that a single-step conversion of methane to methanol would have farreaching economic implications. The direct partial oxidation of methane has been investigated by several authors. Recently published review by Garcia and Lóffler (1984), Foster (1985), Gesser et al. 0888-5885/91/2630-2287$02.50/0
Klier (1986) present a good general
the field. The published results of the homogeneous gas phase reaction seem to be very promising in respect to methanol selectivity, but in most of the studies the yield is still far too low to be of economic interest. Surveying the present literature on the partial oxidation of methane, it appears that systematic studies which focus the influence of the reaction parameters (pressure, temperature, space time, oxygen feed concentration) on the selectivity and the conversion are scarce. With this in mind we have studied systematically the influence of these parameters using a specially designed tubular reactor packed with quartz chips. It will be shown that, under the conditions used, the space time can be shortened considerably compared to earlier investigations. on
Experimental Section Apparatus. All experiments were carried out in a fully
computer controlled apparatus. A process flow diagram of the reactor system is shown in Figure 1. The reactor consisted of an Inconel tube with an inner diameter of 10 mm and length of 150 mm, of which 60 mm were heated by an electric resistance wire (Philips Coax). Into this tube a close-fitting quartz tube was inserted, making the actual internal diameter 8 mm; this gave a heated reactor volume of 3 cm3. To allow comparison with future experiments using catalysts, the free reactor volume was filled with quartz chips (Wisag, Zürich) with a diameter in the range of 1-2 mm. All temperatures were measured with Inconel-clad type K thermocouples (Philips Coax). The reactor temperature was controlled by a computer. For safety reasons the temperatures were measured at various points of the apparatus (Figure 1) and registered with a computer to recognize hazardous situations immediately. The flow rates of the reactant gases (pure CH4 and pure 02) were measured and controlled by mass-flow controllers (Brooks). After these flow controllers, check valves were installed to ensure unidirectional flow. The gases were preheated ©
1991 American Chemical Society
