1854
Ind. Eng. Chem. Res. 1987, 26, 1854-1860
The tentative assumption was made that the high MW lignin in KBL was responsible for multivalent ion retention. Figure 8 indicates that this assumption was not totally justified. When a white liquor (diluted 1:l with water) was run through 50000 MWCO membranes, the “retentate” from this white liquor had some tendency to retain a disproportionate percentage of the multivalents. The white liquor, like its corresponding KBL obtained a t the same mill, was filtered through nylon bags with 100-Fm pores before a UF run. Despite this prefiltration, carbon particulates subsequently settled out from the white liquor. When these particulates were more carefully removed, a smaller percentage of the multivalents were found in the “retentate”. However, even with this doubly filtered white liquor, some residual multivalent retention was observed. For Ca and Mg, about 30% of each was retained as compared to 20% for the Na control. For Fe, about 40% was still retained. The results indicate that some small percentage of Ca and Mg and a larger percentage of Fe may be present as microparticulates or colloids. It is well-known that the greater the valency of an ion and the higher the atomic number of the element involved, the more tightly it is bound to an ion exchanger. For example, Ca2+would be bound to polystyrene sulfonate more strongly than Na+. In the present situation, lignin molecules, including phenolic hydroxyls, would be negatively charged at the high pH (ca. 12.5-13) of the KBLs. This basicity would accentuate complex formation with the positively charged metal ions. Why the higher MW lignin would be a preferential ligand for the multivalent ions as compared to the lower MW lignin found in the permeate is not clear. The di- and trivalent ions may well interact with more than one lignin molecule to form a higher MW aggregate which is then rejected by a given membrane. In addition, that portion of the lignin found in the concentration polarization boundary layer may be partially responsible for the retention. It would be worthwhile to examine the effects of varying degrees of concentration polarization on the retention seen. Evaporator fouling during KBL concentration is primarily due to that portion of Ca bound to organic components in the KBL (Venkatesh and Nguyen, 1985) rather
than to that portion present as insoluble carbonates or slightly soluble sulfates. On the basis of results shown here, it may be tentatively assumed that most of the Ca removed from KBL by UF is bound to high MW lignin, i.e., that is is probably a fouling form of Ca. Tests are needed to examine the relative effects on heat transfer of the feed KBL as compared to its permeate. Conclusions The ultrafiltration of KBL can be used as a means to remove a significant proportion of multivalent ions into the retentate. The element most effectively removed is Mg, but a substantial proportion of Ca is also found in the retentate. Most of the multivalents removed are probably associated with lignin. However, a more than proportionate level of Ca, Mg, and Fe was found even in a white liquor retentate; thus, some level of these elements is probably present as particulates removed directly by the filter. This work was funded by the Department of Energy, Office of Industrial Programs. The technical assistance of Dale Violette is gratefully acknowledged. Registry No. Ca, 7440-70-2; Mg, 7439-95-4; Al, 7429-90-5; Fe, 7439-89-6.
Literature Cited Akred, A. R.; Fane, A. G.; Friend, J. P. In Ultrafiltration Membranes and Applications; Cooper, A. R., Ed.; Plenum: Press, New York, 1980; “Polymer Science and Technology”, Vol. 13, p 353. Cooper, A. R.; Booth, R. G. In Ultrafiltration Membranes and Applications; Cooper, A. R., Ed.; Plenum: New York, 1980; “Polymer Science and Technology”, Vol. 13, p 283. Grace, T. M.; Sachs, D. G.; Grady, H. J. Tappi 1977, 604(4), 122. Hill, M. K. TAPPI Proceedings of the 1985 Pulping Conference, Hollywood, FL, 1985; p 629. Hoffman, M. R.; Yost, E. C.; Eisenreich, S. J.; Maier, W. J. Enuiron. Sci. Technol. 1981, 15(6), 655. Kirbawy, A., Weyerhaeuser Co., unpublished data, 1984. McLellan, J. K.; Rock, C. A. Int. Peat J. 1986, 1 , 1. Venkatesh, V.; Nguyen, X. N. In Chemical Recouery in the Alkaline Pulping Processes; Hough, G., Ed.; TAPPI: Atlanta, 1985; p 45. West Virginia Pulp and Paper Co., Technical Bulletin 107, 1954.
Received for review May 15, 1986 Accepted June 12, 1987
Deactivation Kinetics of Toluene Disproportionation over Hydrogen Mordenite Catalyst Shiba Prasad Bharati and Subhash B h a t i a * Department o f Chemical Engineering, Indian Institute of Technology, Kanpur 208016, India
T h e kinetics and mechanism of deactivation by coking of hydrogen mordenite catalyst for the disproportionation of toluene have been studied along with the kinetics of the main reaction. A statistically best rate expression for the main reaction, developed on the basis of the single-site adsorption of toluene, was determined from the experimental data. Deactivation occurred in parallel with the main reaction where toluene was adsorbed in different ways in the main and deactivation reactions. The deactivation kinetic equation was represented by a heterogeneous model incorporating coke formation by a parallel reaction scheme. Research on toluene disproportionation catalysis to produce benzene and xylene has gained attention in recent years due to the increasing demand for xylene in the petrochemical industry. A large number of patents have
* To whom correspondence should be addressed. 0888-5885/87/2626-1854$01.50/0
appeared which describe a variety of solid acidic catalysts for heterogeneous vaporlliquid-phase reactions for toluene disproportionation. Vapor-phase systems with heterogeneous solid catalysts are economically significant and viable. Commercial processes have been announced by Sinclair/Atlantic Richfield (Oliver and Inouse, 1970) and 0 1987 American Chemical Society
Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 1855 Toray Industries/Universal Oil Products (Ponder, 1979). Many variations of process schemes for transformation of relatively low cost aromatic mixtures to desired high value products, such as p-xylene, are becoming available. The selective preparation of para-substituted benzene derivatives by means of various modified zeolite catalysts has recently been reported (Kaeding, 1981). In particular ZSM-5 class zeolites have been of considerable interest because the channel dimensions are approximately the same as the molecular dimensions of many aromatic molecules. Mordenite- and faujasite-type zeolites have also been used, and it was found that hydrogen mordenite had higher activity compared to hydrogen-Y zeolites (Minachev et al., 1971). Shasidhar and Patwardhan (1981) studied toluene disproportionation reaction and analyzed their kinetic data based on a reversible second-order reaction model, and the deactivation reaction was represented by a simple power law model. Earlier studies on catalyst deactivation contemplated the catalyst activity as dependent only on time, and an exponential relationship between activity and time was reported (Shasidhar and Patwardhan, 1981). These investigations revealed little information about the mechanism of deactivation. The deactivation studies of catalyst particles under diffusional control (Kaang and Levenspiel, 1973) and in fixed bed reactors (Weng, 1975) have demonstrated that the deactivation also depends upon the reaction conditions. Mordenite catalyst, being sensitive to coking because of its porous structure, deactivates fast compared to other zeolite catalysts like zeolite Y. Recently, Corella and Asua proposed (1981) and generalized (1982) theoretically the kinetics of deactivation by coking, which relates activity directly to the deactivation reaction, and obtained the deactivation equation of the form
+
where +(pi, T )is the deactivation function and d = ( m h - l ) / m , m and h beiqg the number of active sites involved in the controlling step of the main reaction and deactivation reaction, respectively. The starting point for obtaining analytically the kinetic equations for deactivation for each mechanism is based on the fact that for a simple reaction, i.e., of the type A R + S, the activity is given by
-
as has been found by Chu (1968), Jodra et al. (19761, Froment (1976), Beckman and Froment (1979), and Froment and Bischoff (1979). L is the total concentration of active sites and CP1 is the concentration of active sites covered by coke. The method has been successfully applied in the studies of catalyst deactivation by coke formation in dehydration of isoamyl alcohol over silica alumina catalyst (Corella and Asua, 1981) and methylcyclohexane dehydrogenation over Pt-Re-A1203 catalyst (Srivastava et al., 1986). The present study has been undertaken to obtain rate and selectivity data for toluene disproportionation using hydrogen mordenite catalyst in a fixed bed isothermal reactor. Hydrogen atmosphere was used to keep the catalysts free from deactivation due to coke formation throughout the reaction. The experimental results were
analyzed on the basis of Langmuir-Hinshelwood kinetics for the main reaction as well as for deactivation of mordenite due to coking reaction. Experimental Section Catalyst. Zeolite hydrogen mordenite (Zeolon 900H) was obtained from M/S Norton Chemicals, England, with silica-alumina ratio of 10 and effective pore diameter of 8.9 A. The nickel-hydrogen mordenite (Ni-HM) catalysts containing nickel metal on hydrogen mordenite (HM) was prepared by ion-exchange technique. The H form of mordenite was subjected to repeated exchange with nickel nitrate solution under reflux conditions, followed by washing with distilled water and drying at 378 K. Aluminum-deficient mordenite was prepared from hydrogen mordenite by leaching aluminum with hydrochloric acid. After leaching, the catalyst samples were washed with distilled water until they were acid free and dried at 378 K. The catalyst was calcined in a muffle furnace at a constant temperature of 773 K for 8 h. The catalyst was reduced in the reactor in flowing hydrogen (40 mL/min) for 3 h at a constant reduction temperature of 673 K. Experimental Setup Toluene disproportionation was carried out in a fixed bed flow reactor at atmospheric pressure. Toluene was passed through a metering pump (Model RP-G 20, range 0-55 mL/min) supplied by Fluid Metering, Instrument Corporation, New York. The reaction was carried out in hydrogen atmosphere to suppress catalyst deactivation due to coke formation. The toluene/H2 ratio was maintained as 1.5. The reaction mixture (toluene-hydrogen) as preheated in a tubular furnace which was controlled within an accuracy of f 5 K by a digital temperature controller using iron constatan thermocouple. The stainless steel reactor of dimensions 1.5 X 2.5 cm was used as a fixed bed reactor. The catalyst bed was supported by ceramic beads below and above it for proper distribution of reactants and heat transfer. The temperature of the catalyst bed was controlled within an accuracy of fl K by an APLAB temperature controller using a chromel-alumel thermocouple placed centrally in the reactor. The main stream of the effluents from the reactor was cooled by circulating ice-cooled water in a stainless steel coil-type condenser. Liquid samples were analyzed over a gas-liquid chromatograph with a flame ionization detector. The components of the sample mixture were separated over a 16-mm X 3-m long stainless steel column packed with 5% Bentone-3,4-diisodecyl phthalate on 6080-mesh Chromosorb W. Kinetic data were taken at different temperatures ranging from 623 to 723 K a t a constant value of liquid hourly space velocity (LHSV). The experimental runs were taken for different LHSV values varying from 10 to 40 (g of catalyst)/(g-mo1.h). In order to study the deactivation kinetics, the experiments were carried out with different times on stream for a given value of temperature and LHSV. Samples were analyzed at different time intervals to check the activity of the catalyst. Experimental data were taken at different values of temperature and LHSV. Results a n d Discussion It is very important that mass-transfer effects should be absent during the reaction in a fixed bed catalytic reactor. Mass-transfer effects, both external, and particle diffusion were checked in the present study before reaction
1856 Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987
I
Toluene / Hp ( m o l e I = 1.5 Catalyst
Hydrogen m o r d e n i t e ( H - M )
+ 723K
25
I
W / F ( g - c a t hr /g.mole I
Figure 1. Toluene conversion vs. time factor.
kinetic data, and the deactivation kinetic study was undertaken. Mass Transfer Effects. The important variables which can affect kinetics of the reaction due to masstransfer resistance are liquid hourly space velocity (LHSV) and catalyst particle size. Experimental runs were carried out to study the effect of both variables. Two different particles sizes of 1.5 mm and finely ground catalyst were used to test pore diffusion limitations. The effect of external diffusion on conversion, defined as moles of toluene reacted per unit moles of toluene fed in the reactor, was studied by conducting two runs with fixed W / F but varying the weight of catalyst. The weight of the catalyst in one run was almost double the weight in another run. The external mass-transfer effects and pore diffusion resistance were found to be absent under experimental conditions. Catalyst particle size of 1.5 mm was used in subsequent experimental runs. Effect of Carrier Gas. To study the effect of carrier gas on toluene disproportionation, nitrogen as well as hydrogen gas was used. Under nitrogen gas, the catalyst deactivation was observed to be much more rapid, as compared to hydrogen gas atmosphere. Mass spectrum studies (Minachev et al., 1971) of desorbed products of the coked mordenite reported that the nature of the coke depends upon gaseous environment. Futher, it has been reported that the coke deposited under nitrogen atmosphere is a higher molecular weight compound that the one formed under hydrogen atmosphere. Carbonium ions were found to be the intermediates for coking and disproportionation reactions (Streitwieser and Reif, 1964). Hydrogen mordenite, which exhibits very strong acid sites, was found to be capable of activating hydrogen in such a ways as to enable it to reduce the concentration of adsorbed carbonium ion species on the catalyst surface and thereby minimize the coke formation. The effect of hydrogen gas on paraffinic transformation using acidic catalysts has been reported in the literature (Hogeveen and Garbeek, 1969; Streitwieser and Reif, 1964). In hydrogen mordenite catalyst, it is also known that molecular hydrogen can be activated in a superacid medium, by strong acid sites, which can remove coke from acid sites. Thus, deactivation of hydrogen mordenite under hydrogen atmosphere will be lower as compared to the inert atmosphere of nitrogen. Reaction Kinetics and Modeling. Kinetics of toluene disproportionation was studied in an integral reactor. The catalyst was stable for 1h, giving constant conversion. It is assumed that catalyst deactivation did not take place during this period of time and kinetic data were taken in
w/F
(
g catalyst - h r / g mole 1
Figure 2. Test for reversible second-order model.
the absence of deactivation. The experiments were conducted with a constant-value toluene/hydrogen mole ratio of 1.5 at five different temperatures (623, 648, 673, 698, and 723 K) in the W / F range of 10-40 g of catalyst.h/ (g-mol of toluene). The results are plotted in Figure 1. The numerical values of the rate of reaction at zero time, (-rT)o, were obtained by a second-order polynomial fit of the curves (X& vs. W/F followed by analytical differentiation. Homogeneous Model. A homogeneous model based on second-order reversible reaction was proposed for toluene disproportionation (Shasidhar and Patwardhan, 1981) k
P(to1uene)
benzene
PT
PB
+ xylene
(3)
PX
The rate of reaction, expressed as gram moles of toluene disappeared per hour per gram of catalyst, is represented as
dF
rate = --d W = k(PT2- PBPx/K)
(4)
Integrating the above expression will give
+ B - (l/K)’/’ 2Ax + B + (l/K)liz 2Ax
1-
where W = weight of catalyst, g; F = molal flow rate of toluene, g-mol/h; K = equilibrium constant; k = reaction rate constant, g-mol/(h.atm2.g of catalyst); x = toluene conversion; PTo= partial pressure of toluene in the feed stream; A = 1 - 1 / 4 K and B = -2.0. The kinetic data were analyzed by using eq 5 . The test for the validity of eq 5 is shown in Figure 2. These data show a linear relationship, confirming the validity of the reversible second-order rate model. The rate constants at different temperatures are calculated from the slope of linear plot shown in Figure 2. Heterogeneous Model. The disproportionation of toluene to benzene and xylene can be represented on the catalyst surface as 2(toluene) == benzene + xylene (6) T B X The following steps are considered to take place succes-
Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 1857 Table I. Isothermal Rate Equations model
rate-controlling step
SA
single-site mechanism adsorption of toluene
rate equation Catalyst : Hydrogen mordenite Toluene /Hydrogen = 1.5 (-rT)o
=
KTPBPX
+ KBPB + K x P x
l+-
KPPT
DS
dual-site mechanism surface reaction
=
k i K T ( P T 2 - PBPx/K) (1 + KTPT + K$B + KxPx)'
sively a t the surface when using the Langmuir-Hinshelwood kinetic approach. (1) Toluene is adsorbed on the surface of the catalyst. (2) There is surface reaction between adsorbed reactants. (3) There is desorption of the products from the surface. Various plausible kinetic models were derived based on ita reaction mechanism, such as (a) rate of adsorption controlling (single site), (b) single-site surface reaction controlling, and (c) dual-site surface reaction controlling. The isothermal rate equations, basic types of which are shown in Table I, thus resulted in 16 different mathematical forms which were confronted with the experimental data. A linear-regression algorithm utilizing the NewtonGauss technique (Kiovsky and Goyette, 1978) was used to obtain a mathematical fit for various rate models. The model discrimination was done based upon the requirement that kinetic and adsorption constants had to be positive. After the improper models were cancelled, the selected models were tested by nonlinear regression analysis. The rate models with all positive constants which had to be retained after the isothermal regressions are given below. model
SA-1
rate-controlling step
single-site adsorption
rate equation
(-+)a
Figure 3. Arrhenius plots for the rate constants and adsorption equilibrium constant of the rate equation. Table 11. Values of t h e Kinetic Constants T,K k , , g.-mol.h-'.g-'.atm-' K,, atm-' 623 0.040 f 0.00012 13.67 f 0.410 648 0.0059 f 0.00029 11.30 f 0.565 673 0.0079 f 0.00024 8.02 f 0.241 698 0.0125 k 0.00013 6.09 f 0.610 723 0.016 f 0.00032 5.91 f 0.236
The kinetic parameters evaluated from heterogeneous models are shown in Table 11. Dependency of various kinetic parameters on temperature was determined by
k = 33.3 ex,( -
=
y)
k, = 206.8 ex.(-?) SA-2
single-site adsorption
(-rT)o
KB = 3238 exp(
= +
KBPB
The above models were derived from the rate expression
where adsorption of toluene is the rate-controlling step, i.e., T + S + TS adsorption
TS + T = BS + X surface reaction (Eley-Rideal mechanism)
BS + B 2T + B
+S
+X
desorption
overall reaction
3779 7)
where k = homogeneous rate constant, g-mol/(g of catalyst.h.atm2); k l = heterogeneous rate constant, g-mol/ (g of catalyst-h-atm);and KB = adsorption equilibrium constant, atm-'. The apparent activation energy for toluene disproportionation reaction was 12.4 kcal/mol in the case of the heterogeneous model and 11.3 kcal/mol for the homogeneous model. These values were in agreement with the reported value of 14.5 kcal/mol in the literature (Shasidhar and Patwardhan, 1981). Kinetics of Deactivation. In the case of hydrocarbon reactions over mordenite catalyst, it has been observed that high catalytic activity of the zeolite was reduced more or less rapidly with time on stream. One possible reason for deactivation could be due to coking, where the coke can deactive the catalyst by covering active sites and by blocking pores. Some of the coke molecules are large enough to block a pore. Therefore site coverage is accounted for by considering an activity function for (a) main reaction and (b) coking. In order to determine whether the deactivation is in parallel or series with the main reaction, preliminary ex-
1858 Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 Time .hr-
201
Toluene T i m e , hr
15'10
c Ln L
D
-
$ 0
-
c
c
C
,o 25-
Time hr
0
a
2oc
Tolume _- Hydrogen
Toluene
:6
Temp
48 K Time ,hr
1.5
1
0
5
10
15 20 25 30 W / F ( g - c o t h r /g m o l e )
35
W/F(g-cal hr/gmoll
4C
Figure 4. Conversion against space time for different process times a t different temperatures.
periments were performed with 10% benzene (one of the products) in the feed. The results show that addition of benzene had no appreciable effect on the reaction rate. These experiments show an indication of parallel deactivation. Assuming that parallel deactivation takes place on the catalytic surface, the reaction scheme can be represented as 2(toluene) F+ benzene + xylene (11) toluene
-
coke
(-rT),
+T),O
at I T or Pi)
1.0
L
Toluene 215 Hydrogen a 6d8K A 673 0
698 723
(12)
In order to determine the kinetic equation for the deactivation reaction, the deactivation rate data were taken from the experiments carried out in an integral reactor for 3 h with the following operating conditions: toluene/H, ratio = 1.5, temperature range, 648-723 K; WIF = 10-40. These data are represented in Figures 4 and 5. From the X vs. WIF curves at different operation times (Figures 4 and 51, the reaction rate, which is the slope of the curve, was calculated for a particular space time, (W/@ = 40, at different operation times: (-rT)o, (-rT)l, (-rT),, etc. Then the activity at any time t was determined as the ratio between the reaction rate at any time t , (-rT),, and the reaction rate at zero time, (-rT),o,. with both rates measured at the same conversion or partial pressure of the reaction mixture utilizing the relation (Corella and Asua, 1982) a, = -
Figure 5 . Conversion against space time for different process times a t different temperatures.
(13)
The values of the activities at different times are shown in Figure 6. The deactivation rates (-da/dt) were evaluated again by the usual second-order polynomial fitting as well as analytical differentiation. The deactivation function values were calculated from eq 1 and are listed in Table 111. The possible isothermal rate equations for the deactivation reaction (eq 12) were derived based on single-site
0
I
!
!
1
I
5
10
15 Time l h r I
20
2 5
30
Figure 6. Activity against process time.
and dual-site mechanisms by the same method as described by Corella and Asua (1982) and are represented in Table IV. All these derivations were based on the assumption that the controlling step was the formation of the coke precursor (step 2). To evaluate the parameters of the deactivation rate equation, the value of KB (eq 10) for the main reaction was also employed. The model which described the experimental results most satisfactorily is (model DE-3) k&T* 2 P ~ 2 -da- a2 dt (1 + KBPB + KT*PT)'
(14)
The model was based on the assumption that two adjacent adsorbed toluene molecules are involved in the formation of the coke precursor P,lz which ultimately leads to coke through several equilibrium steps.
Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 1859 Table 111. Deactivation Function Values: Initial Rate (-rT)O, Rate at Any Time (-rT),and Corresponding Partial Pressure Value at W / F = 40, R = 1.5 t, h 103(-r~)~ 103(-rT) a PT PB T,K -da/dt *(PI,Ir? 4.50 4.50 1.0 0.0 0.45 0.075 723 1.26 1.260 0.5 1.0 1.5 2.0 2.5 3.0
6.79 8.40 10.00 10.76 14.10 13.14
3.60 2.27 1.60 1.40 1.20 0.94
0.53 0.27 0.16 0.13 0.085 0.07
0.465 0.510 0.540 0.564 0.57 0.576
0.068 0.045 0.030 0.018 0.015 0.012
0.0 0.5 1.0 1.5 2.0 2.5 3.0
4.0 4.83 5.45 5.55 6.25 6.97 6.80
4.0 2.9 1.7 1.0 0.78 0.67 0.51
1.0 0.960 0.31 0.18 0.12 0.096 0.075
0.48 0.498 0.534 0.522 0.570 0.576 0.582
0.060 0.051 0.033 0.024 0.015 0.012 0.009
698
0.0 0.5 1.0 1.5 2.0 2.5 3.0
3.75 3.25 3.67 6.00 7.46 8.50 8.47
3.75 2.18 1.43 1.38 1.22 1.85 0.72
1.0 0.67 0.39 0.23 0.15 0.10 0.085
0.51 0.519 0.540 0.555 0.573 0.579 0.582
0.045 0.040 0.030 0.023 0.014 0.010 0.009
673
0.0 0.5 1.0 1.5 2.0 2.5 3.0
2.67 2.97 3.46 3.88 3.98 4.12 4.50
2.67 2.17 1.63 1.164 0.80 0.49 0.45
1.0 0.73 0.47 0.30 0.20 0.12 0.10
0.525 0.531 0.552 0.569 0.576 0.582 0.585
0.038 0.035 0.024 0.018 0.012 0.009 0.008
648
-
step 2: nM(g) + hM*l P1lh step 3: (Plln)(P21h)* (P31h)--coke m
h
DE-1
1
1
DE-2
1
1
n 0
,
DE-3
1
2
0
DE-4
1
2
1
rate equations
da kdkT*PT dt 1 + KBPB + KT*PTa da dt
da dt
--=
kdkT*PT2
1
+ K B P B + KT*PTa kdKT*'PT'
(1 + KBPB + KT*PT)'
a2
kdK~**p~~ da a2 dt (I + KBPB + KT*PT)'
1.873 f 0.1315 1.848 f 0.13 1.629 f 0.202 1.545 f 0.155
0.072 f 0.0053 0.111 h 0.0091 0.091 f 0.011 0.158 f 0.016
The parameter values corresponding to eq 14 are listed in Table V. The relations between the constants of the deactivation rate equation and temperatures are kd
= 2075 eXp(
0.941 0.57 0.33 0.20 0.14 0.097 0.08 0.88 0.578 0.456 0.339 0.228 0.117 0.117
0.941 1.269 2.169 3.780 6.222 9.700 11.073 0.880 1.085 2.064 3.767 5.700 8.125 11.700
Acknowledgment
Table V. Kinetic Constants for the Deactivation Reaction T,K k d , h-' KT*, atm-' 648 673 698 723
2.136 3.566 5.469 5.917 8.304 9.184 0.969 1.611 3.486 8.462 7.639 9.765 12.444
Conclusion The rate equation for the toluene disproportionation along with the kinetics and mechanism of deactivation of the hydrogen mordenite catalyst has been determined. The kinetic data were most satisfactorily correlated by homogeneous and heterogeneous models. The activation energies calculated from both models were comparable. Deactivation of hydrogen mordenite catalyst due to coking was represented by a heterogeneous model in which the coke formation reaction was assumed to be parallel with the main toluene disproportionation reaction.
Table IV. Isothermal Deactivation Rate Equations Assuming Step 2 as Rate Controlling step 1: M + 1 + M*l
model
0.60 0.26 0.140 0.10 0.06 0.045 0.969 0.58 0.335 0.28 0.11 0.09 0.07
F) T)
KT* = 49 exp( 4924
The activation energy for the deactivation was 4.70 kcal/mol which is less than that of the main reaction, indicating fast deactivation of the catalyst. Also, the deactivation reaction rate is more sensitive to temperature compared to the main reaction.
We gratefully acknowledge M/S Norton Chemicals, England, for supplying the free samples of hydrogen mordenite catalyst used in the present study. Nomenclature a = catalyst activity, integral reactor a = catalyst activity at time j d,, = concentration of active site covered by coke C, = concentration of component i d = deactivation order F = feed rate of toluene, mol/h h = number of active sites involved in the controlling step of the deactivation reaction K = thermodynamic equilibrium constant for the main reaction KT, KB,Kx = equilibrium adsorption constants for toluene, benzene, and xylene, respectively, in the main reaction, atm-l KT* = equilibrium adsorption constant for toluene to yield coke precursors kl, k = rate constant of the main reaction for heterogeneous and homogeneous models, g-mol/ (g of catalyst-hsatm) k d = rate constant of the deactivation reaction, h-' 1 = active site m = number of active sites involved in the controlling step of the main reaction n = number of molecules of T in gas phase which react with adsorbed T to give the coke precursor
Ind. Eng. Chem. Res 1987,26, 1860-1864
1860
PI&,P,lh = coke precursor P212,P,lh, P312,P31h = different form of coke in the coking sequence PT,PB,Px.= partial pressures of toluene, benzene, and xylene, respectively, atm P, = partial pressure of component i, atm R = mole ratio of toluene to hydrogen -rT = rate of reaction of T, g-mol/(g of catalyst-h) +T)O = rate of reaction of T at zero time, g-mol/(g of cata1yst.h) T = absolute temperature, K t = time, h T, B, X = toluene, benzene, and xylene, respectively T1 = toluene adsorbed in the main reaction T * l = toluene adsorbed in reaction different from the main reaction W = weight of catalyst in reactor, g X = conversion of toluene ~~
~
Greek Symbols
+ ( p L ,7‘) = deactivation function, general +(pT, 7‘) = deactivation function, T system Registry No. C6H,CH3, 108-88-3; C&, 71-43-2; CH3C6H4CH3, 1330-20-7; C, 7440-44-0.
Literature Cited Beckman, J. W.; Froment, G. F. Ind. Eng. Chem. Fundam. 1979,18, 245.
Chu, C. Ind. Eng. Chem. Fundam. 1968,7,509. Corella, J.; Asua, J. M. Can. J. Chem. Eng. 1981,59,647. Corella, J.; Asua, J. M. Ind. Eng. Chem. Process Des. Dev. 1982,30, 293. Froment, G. F. “Proceedings of the 6th International Congress on Catalysis”, London, 1976. Froment, G. F.; Bischoff, K. B. Chemical Reactor Analysis and Design; Wiley: New York, 1979. Hogeveen, H.; Garbeek, C. J. Recl. Trans. Chim. Pays-Bas 1969,88, 719. Jorda, L. G.; Romero, A,; Corella, J. An. Quim. 1976,72,823. Kaang, S. J.; Levenspiel, 0. Ind. Eng. Chem. Fundam. 1973,12,185. Kaeding, W.W. J . Catal. 1981,69,392. Kiovsky, J. R.; Goyettee, W. J. J. Catal. 1978,52,25. Minachev, Kh.; Garanin, V. I.; Isakova, T.; Kharlamov, V. V. Molecular Sieve Zeolites-II; Advances in Chemistry Series 102; American Chemical Society: Washington, D.C., 1971; p 441. Oliver, E. D.; Inouse, T. Aromatics B T X ; Stanford Research Institute: Stanford, CA, Hand Book No. 30A, 1970. Ponder, T. Hydrocarbon Process. 1979,141. Shasidhar, B. S.; Patwardhan, S. R. Ind. Eng. Chem. Prod. Res. Deu. 1981,20, 102. Srivastava, R. D.; Bhowmick, M.; Pal, A. K. Ind. Eng. Chem. Process Des. Deu. 1986, 25, 236. Streitwieser, A.; Reif, L. J. Am. Chem. SOC.1964,92,3831. Weng, H. S. Chem. Eng. Sci. 1975,30, 1341.
Received for review July 18, 1986 Revised manuscript received May 27, 1987 Accepted June 8, 1987
Deactivation of Y-Zeolite Catalysts by Coking in Methylnaphthalene Isomerization Lucio Forni* Dipartimento di Chimica Fisica ed Elettrochimica, Universitci di Milano, I-20133 Milano, Italy
Vincenzo Solinas and Roberto Monaci Dipartimento di Scienze Chimiche, Uniuersitci di Cagliari, I-09100 Cagliari, Italy
T h e progressive deactivation of variously decationated Y-zeolite catalysts in the title reaction was studied at atmospheric pressure and 543-663 K. T h e activity decay with time on-stream is satisfactorily expressed by a very simple exponential equation, whose parameters depend on the degree of decationation of the zeolite and on temperature. Coke is formed of insoluble, highly condensed material, which can be easily burnt off, restoring the behavior of fresh catalyst. Coke formation takes place in two successive periods, with different rates, both following the Voohries’ equation. Only t h e slow deactivation due t o weaker centers is a n activated process. T h e faster coking, due to stronger acid sites, is practically independent of temperature. In a previous work (Solinas et al., 1984), it was found that Y-zeolite, totally or partially protonated, is a very active and selective catalyst for the isomerization of 1methylnaphthalene (1MN) to 2-methylnaphthalene (2MN), the latter being a well-known useful intermediate for several applications in the field of fine chemicals. However, due to the presence of relatively strong acid sites at any ion-exchange degree, a quite rapid decay of activity was also observed, due to fouling by carbonaceous deposits. Coke deposition is the main reason for catalyst deactivation in organic reactions and many papers have been published on this subject in recent years (Beekman and Froment, 1979; Froment, 1980; Corella and Asila, 1982; Wukasch and Rase, 1982; Blackmond et al., 1982; Lin et al., 1983; Bilbao et al., 1985; Derouane, 1985; Barbier, 1986). However, the word “coke”, as referred to in the literature, cannot find a simple definition. It is generally accepted that it is not a single component, but it includes all the carbonaceous material remaining on the catalyst after the reaction. Moreover, it continuously evolves,
during the reaction, to higher and higher condensation grades. It is also usually accepted that the activity of the zeolite catalysts is affected by coke deposition in different ways, e.g., by site coverage or pore ,blockage (Derouane, 1985). The deactivation is faster in the second case, the pore blockage preventing the access of reactants to a higher amount of active sites. In the present study, the deactivation of Y-zeolites, due to coke deposition during the mentioned isomerization reaction, was examined as a function of temperature, time on-stream, and surface acidity of the catalyst, looking for correlations useful in the development of such a process. Experimental Section Catalysts. Y-Zeolite catalysts were prepared from commercial crystalline powder cake (LZY-52,from Union Carbide) by ion exchange in the usual way (Bolton, 1976). After ion exchange, drying, and calcination (12 h at 773 K), the powder was pressed (maximum pressure of 200 MPa) in ca. 1 mm thick wafers, which were then gently
0888-588518712626-1860$01.50/0 0 1987 American Chemical Society
