Ind. Eng. Chem. Process Des. Dev. 1983, 22, 639-645
.
0.50
I
I
I
I
parison of the Chao-Seader and RKJZ predictions with experimental data on Wyoming coal liquids reported here has confirmed the conclusions reached for Illinois coal liquids (Wilson et al., 1981). In predicting the volatility of coal liquids, RKJZ with either MB or the new vapor pressure method is generally equivalent to Chao-Seader with the new method, but is clearly superior to the Chao-Seader/MB combination. In predicting the K values of light gases, replacement of MB by the new method improves the RKJZ prediction significantly but has no effect on Chao-Seader. More data are needed to better establish the temperature dependence of the H2and methane K values. The experimental data on both Illinois and Wyoming coal liquids indicate that the coal source apparently has little effect on the volatility of VLE behavior of coal liquids, provided the coal liquids have a similar boiling-point distribution.
I
ExwrimmUi (P = 1950 p i i d
0,45
- 700 OF
0.40
-
0
775 860 OF 'F
+
0 Ill. COIl
Liq. I
111. Coil Liq. Ii
Pndictloni
0.30
-
0.25
-
0.20
-
0.15
-
639
Acknowledgment James R. Freeman and Richard S. Owens of Wiltec Research carried out the experimental measurements. 00
0.25
0.50
075
100
125
150
Registry No. Hydrogen, 1333-74-0;methane, 74-82-8.
175
WEIGHT % H2 CHARGED
Literature Cited
Figure 5. Effect of H2on vaporization of coal liquids.
Hwang, S. C.; Tsonopoulos, C.: Cunningham, J. R.; Wilson, G. M. I n d . Eng. Chem. PrOcessDes. D e v . 1982, 21, 127. Vlck, G. K.: Epperly, W. R. Science 1982, 217(4557), 311. Wilson. G. M.: Johnston, R. H.; Hwang, S. C.; Tsonopoulos, C. I n d . Eng. Chem. Process D e s . D e v . 1981, 20, 94.
apparently has little effect on the volatility or VLE behavior of coal liquids, provided the liquids have a similar boiling-point distribution. Conclusions The study of the volatility of EDS coal liquids has been extended to liquids produced from Wyoming coal. Com-
Received for review October 18, 1982 Accepted March 24, 1983
Cyclohexane Dehydrogenation for Thermochemical Energy Conversion George 8. DeLancey, Suphan Kovenklioglu; Arthur B. Ritter,+ and James C. Schneidert Stevens Institute of TechnoicqY, Department of Chemistry and Chemical Engineering, Hoboken, New Jersey 07030
A study of cyclohexane dehydrogenation to benzene and hydrogen at atmospheric pressure and in the temperature range of 533-589 K was carried out with the objective of obtaining rate data to be used in the design and evaluation of a reactor system for energy collection and conversion in a thennochemical cycle. An internal recirculation reactor was employed. External gradients were observed at higher reaction rates. A reaction mechanism where the rate-determinlng step is the A-Q shllt of the adsorbed cyclohexene was found to be more suitable in correlating the rate data than a mechanism where the dissociative adsorption step is rate controlling. Side reactions which were found to be minor over the range of experimental conditions were attributable to impurities in the cyclohexane feedstock. No loss in catalyst activity was observed.
Introduction The economic and efficient storage of thermal energy is intimately linked to the solution of the critical international problem of diminishing availability of traditional energy ~~~s and the utilization of new ones, such s o h
and nuclear, for meeting the daily demands of industrial and domestic users. One attractive alternative is thermochemical storage systems (Mar and Bramlette, 1978). These systems are based upon the controlled utilization of highly energetic chemical reactions which are reversible over the temperature range of interest. The endothermic reaction is driven by the energy source, which defines the upper limit of the temperature range, and the thermal energy is recovered from the exothermic reaction at the energy sink, which defines the lower limit of the temper-
'University of Medicine and Dent&,vin New Jersey, Newark, NJ.
* Lummus Company, Bloomfield, NJ.
0198-4305/83/1122-0639$0 1.50/0
0
1983 American Chemical Society
640
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
ature range. For catalytic reactions, or for reactions where the reactants and products are physically separable, no significant degradation of the absorbed energy occurs during storage, and the reactants and products may be stored at ambient temperatures. In addition to the fact that the destruction and reformulation of chemical bonds can be considerably more energetic than sensible and latent heat effects, which are also viable storage mechanisms, the thermal stability of the energy stored at ambient temperatures is one of the desirable features of thermochemical storage systems. The complete thermochemical system may be located at one site or the reactants and products may be transported between the collection and utilization sites with some or all of the storage in the transportation network. This latter application has been referred to as the chemical heat pipe (Vakil and Kosky, 1976). In either event, thermochemical systems can be operated in closed, nonpolluting cycles. The dehydrogenation of cyclohexane to benzene is a promising candidate by which the thermochemical storage concept can be demonstrated on a commercial scale.
This reaction was selected for a number of interrelated reasons. The enthalpy change is large and the reaction is reversible at atmospheric pressure in the temperature range of 478-589 K. Light water reactors and focused solar collectors are possible energy sources for the endothermic reactor while saturated steam from the exothermic reactor may be used for electrical energy generation. Cyclohexane and benzene may be stored as liquids under ambient conditions while hydrogen may be stored as a compressed gas or in a hydride fotm (Johnson and Reilly, 1978). It should be pointed out, however, that the disadvantage with this process, as with any thermal process where the difference in temperature at which the energy is stored and utilized is not large, is the thermodynamic inefficiency. The technology for the benzene hydrogenation step (energy recovery) is well known and readily available on a commercial scale (Sittig, 1967; Thomas, 1970). For the dehydrogenation reaction, there is a scarcity of design data in the literature in the intermediate temperature range (478-589 K) at atmospheric pressure. Previous studies have generally been directed to the high pressures and temperatures employed in commercial reformers (Haensel et al., 1964; Smith and Prater, 1967) or to the limited conditions required for studies of the catalytic surface (Gland et al., 1975). It is within the context of thermochemical energy storage that the current study of cyclohexane dehydrogenation is undertaken. The study was carried out at atmospheric pressure in the 533-589 K temperature range under stoichiometric reaction conditions on a commercial platinum catalyst with the objective of obtaining data which could be used in the design and evaluation of a reactor system for energy collection and conversion in a thermochemical cycle. Also, the existence of side reactions and the stability of the catalyst were important questions in the study. Experimental System A flow schematic of the experimental facility is shown in Figure 1. Liquid cyclohexane was fed to the reactor through a syringe by a constant infusion pump and was partially vaporized in the inlet tubing, which was electrically heated. The vaporization as well as partial conversion to hydrogen and benzene was completed in the reaction vessel. The effluent from the reactor passed
lli
&
NEEDLE VIILI'E
@
PRESSURE C A W '
0
I'iEPllOCOOPLE
Figure 1. Experimental system.
through heated tubing to a heated gas sampling valve. The valve extracted samples for analysis at programmed intervals. The time required for the chromatograph to respond to changes in reactor conditions was minimized by keeping the diameter and length of tubing between the reactor and sampling valve as small as possible. Under conditions of high product flow rate, a high pressure drop through the sample loop was relieved by partially bypassing the sampling valve. The chromatograph is a Hewlett-Packard Model 5831 A, a digital unit with dual flame ionization detectors. The column consisted of a 6.35 mm 0.d. stainless steel tube, 1.8 m in length, packed with 15% Carbowax 1540 on 60/80 mesh Chromosorb P. The stream leaving the sampling valve passed through heated tubing and entered the condenser where the liquid products were condensed and separated. Chilled water at 275 K was recirculated through a refrigeration bath. The flow rate of the remaining vapor stream which exits to the exhaust hood was measured by a rotameter. A micrometer type needle valve located at the exit controlled the system pressure. The reactor pressure was measured by a Bourdon type gauge connected to the reactor outlet tubing. Nitrogen or hydrogen was admitted to the reactor for purging and catalyst activation at controlled pressures and rates. Pressure was also measured by a Bourdon gauge downstream of the outlet rotameter. An internal recirculation reactor (Berty et al., 1969) was employed, and a commercial reforming catalyst of platinum supported on alumina (Engelhard Industries RD150-C)was used. The particles were of uniform diameter, but they varied in length. The size distribution data, shown in Figure 2, were obtained by first weighing the particles and then calculating their length on the basis of their diameter and density. To ensure a uniform flow distribution, a catalyst charge of 40 g was interspersed throughout a bed of 3 mm diameter Pyrex glass beads. Inactive layers of both 6 mm and 3 mm diameter beads were placed at the bed entrance and exit. Temperature was measured at three points within the reactor vessel by iron-constantan thermocouples encased in 1.59-mm stainless steel sheaths. One thermocouple was located in the upper turning cavity of the reactor and the other two were located inside the catalyst bed. The thermocouples which were located inside the bed were inserted from the bottom of the reactor at heights of 10 mm above the support grating and 10 mm below the top
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983 641
Table I. Rate Data
12.0
(PT = 1.02 atm)
Cylinders Diameter
=
1.59 m
10.0
-
6 8 0
s m
2
-
6.0
p z w
$ 4.0 Y
e 2.0
bI I
0.04
0.08
0.12
0.16
LENGTH O F P I R T I C L I . t n
0.20
0.24
0.28
Figure 2. Distribution of particle sizes.
of the head. These regions contained only glass beads. Experimental Procedure The system was regularly tested for leaks under a nitrogen pressure of 1.41 atm. The pressure loss criteria were 2.8 X lo4 atm/s in the reactor and 1.9 X lob atm/s in the remainder of the system, which was of considerably smaller volume than the reactor. The system was never operated until the leak criteria were met. Once onstream, the flow rate of the outlet gas was monitored to ensure that the measured flow was equivalent to the expected rate for the prevailing reactor conditions. After installing a new batch of catalyst, or whenever there was a possibility that oxygen had entered the reactor, the catalyst was reduced with hydrogen at 755 K. To begin activation, the reactor was first purged with nitrogen and then with hydrogen. The temperature was raised to 755 K over a period of approximately 5 h and held there for 1h while the reactor was continually purged with hydrogen. Because there is a net increase in the number of molecules in the dehydrogenation reaction, product must be continually withdrawn to maintain the reactor pressure during the startup period, even without feed. Steady state was assumed to have occurred when consecutive samples revealed a trendless variation of less than 0.5 area %. A standard set of operating conditions was established as a means for determining the relative catalyst activity. The activity check was conducted after each catalyst activation and after obtaining each set of data. There was no indication of any loss in catalyst activity. In addition, the system was operated without catalyst under otherwise normal conditions which indicated that the graphite or other surfaces do not act as catalysts. Internal a n d External Gradients The data presented in Table I are free of internal and external gradients. To test for the absence of internal concentration and temperature gradients the following criterion was used (Butt, 1980)
E = 1 f 0.05 if where
TPRp2
1
-< -
ca, 1-rP
r x lo7 g-mol/ X (s g of cat.) T = 533 K 0.420 0.401 0.400 0.763 0.350 1.67 0.315 3.01 0.251 4.79
T = 555 K 0.716 0.674 0.635 0.564
0.341 0.644 1.21 2.69
r~ lo7 g-mol/
X
(s g of cat.) T=575K 0.870 0.414 0.764 0.730 0.690 1.32 0.600 2.86 T = 589 K 0.925 0.440 0.835 0.798 1.41 0.738 2.90 0.610
This criterion is for a first-order reaction in a spherical catalyst. For the cylindrical pellets used here, R, was approximated as 3a where a is defined as the ratio of the apparent volume of the pellet to the external surface area. For a length of 0.28 cm representing the longest catalyst, the above criterion is satisfied for all of the reported data points. Experiments with varying fan speeds were performed to test for external gradients. It was found that at high reaction rates external gradients were present as the reaction rate continued to increase with increasing fan speed. These data points are not reported. All the data presented in Table I are free of external diffusion limitations. We have also noted that at 533 K the conversion for one data point near equilibrium exceeded the thermodynamic equilibrium value by approximately 1%. The cause is believed to be experimental error. This data point is not reported in Table I. The onset of external gradients at high reaction rates is not surprising since the data were collected at atmosphere pressure. The Berty reactor is generally operated over 100 psig where the mass flow rates are high enough to obviate external gradients. For reactions at low pressures it should be possible to eliminate external gradients by operating the reactor at a high partial pressure of an inert diluent such as argon so that the partial pressures of reactants and products can remain the same as at the low pressure originally specified for the study. The problem, however, is that one may introduce internal pore diffusion gradients (Cropley, 1981). This is because the bulk diffusivity is inversely proportional to pr'essure but the Knudsen diffusivity is independent of pressure. As pressure increases, the effective diffusivity, and hence the catalyst effectiveness would decrease unless the diffusion is primarily in the Knudsen regime. Cyclohexane Dehydrogenation Kinetics The simplest expression which has been reported to reasonably correlate cyclohexane dehydrogenation data at 40 atm and 780 K (R. B. Smith, 1959) is given by
Equation 1is consistent with the dissociative adsorption mechanism proposed by Ritchie and Nixon (1968) if the desorption equilibrium constants are large and the dissociative adsorption step is rate controlling, in which event k is the velocity constant. Here cyclohexane goes through the a-,ab- and .rr-adsorbedolefili followed by the stepwise removal of the neutral hydrogen atoms and extension of the ?r system (Germain, 1969). Similarly, Tetenyi (1974) at 523-573 K found the cyclohexene and cyclohexadiene
642
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
Scheme I
Table 11. Thermodynamic Results (PT = 1.02 atm; Pure C,H,, Feed) AHR" I
1. c + I = CZ(U) 2. C l ( u ) = DZ(n) t H,
AGRO,
T, K
kcal/g-mol
kcal/g-mol
Xeq
533 555 575 589
51.94 52.10 52.22 52.32
2.075 -0.081 -1.867 -3.184
0.428 0.753 0.924 0.972
3. D l ( n ) = Dl(a) 4. DZ(u) = EZ(n) t H,
5. El(n) = EZ(u) 6. E l ( u ) = H ( n ) + H,
dehydrogenation rates higher than cyclohexane dehydrogenation and concluded that cyclohexane adsorption must be rate determining. Gland et al. (1975) propose that the slow step in cyclohexane dehydrogenation on the Pt (111) surface is the dehydrogenation of cyclohexene, an intermediate formed at the catalytic surface. Cyclohexene has been observed in the vapor phase at high pressures and temperatures. Haensel et al. (1964), who studied the reaction at 793 K and 20 atm conclude that cyclohexene is an intermediate. Smith and Prater (1967), who studied the reaction at 671 K and 25 atm show that the sequence of reactions observed in the apparent kinetics could differ from the true sequence of surface reactions if the adsorption step is rate controlling. They point out that Haensel et al.'s (1964) conclusion that cyclohexene is an intermediate is not a unique consequence of their observation. At lower pressures and temperatures various investigators have found that the amount of secondary products in the vapor phase is negligible (Usov et al., 1974, Germain, 1969; Bridges and Houghton, 1959). More recently, Ruiz-Vizcaya et al. (1978) have done a theoretical study of the cyclohexane dehydrogenation mechanism by studying the charge distribution of the chemisorbed system. They propose a mechanism which consists of u-adsorption followed by a series of steps where cyclohexane goes through cyclohexene and cyclohexadiene intermediates with H2 elimination. Here, the rate-determining step is the T-u shift of the adsorbed cyclohexene. The implication of this mechanism is the product inhibition by hydrogen in the forward rate as opposed to the first order dependence on cyclohexane predicted by the dissociative adsorption mechanism. This reaction mechanism can be summarized as a succession of the following steps shown in Scheme I (C, cyclohexane; D, cyclohexene; E, cyclohexadiene; B, benzene; 1, catalytic site). The rate expression with step 3 rate determining and all other steps assumed to be in equilibrium is
( 2 'e)/( + +
r = k3KlK2CT[
1
-
KIPc
7 . BZ(n)= B
+1
adsorption H, desorption with simultaneous U-n shift n-u shift; rate-determining step H, desorption with simultaneous U-n shift n-u shift H, desorption with simultaneous U-T shift benzene desorption
where K p = K1K2K&K&& constant in step 3.
and k3 is the forward rate
Experimental Data and Analysis The rate data are presented in Table I. The reproducibility of the results over the three experiments which were repeated ranged from 1 to 5% with an average of 2.5%. Due to the complexity of the rate expression given by eq 2, a simpler form was sought to correlate the rate data. The simplest form of eq 2 is (3) where k' = ka1K2CT,since the thermodynamic limitations on the rate must be reflected. The hydragenation direction is favorable at 478 K and the dehydrogenation direction is favorable at 589 K. The pertinent thermodynamic variables, summarized in Table 11, were calculated without fugacity corrections at atmospheric pressure. Equations 1 and 3 were tested against the data. The results are given in Table 111. The rate constants were determined from linear regression where the combination of terms on the rightrhand side of eq 1 and 3 which appear in parentheses were treated as the independent variable. The thermodynamic requirement is that the regression line pass through the origin. I t should be noted that the regression analysis which minimizes the sum of squares of the differences between the experimental and predicted rates weighs the data points near equilibrium less than those far from equilibrium. From the comparison of relative error (or s u m of squares of the residuals) one observes that eq 3 correlates the rate data somewhat better than eq 1 for the temperature range of 555-589 K. On the other hand, at 533 K, eq 1 gives a better correlation. Also the rate constants for both models do not perceptibly change above 555 K. The increase in the reaction rate above 555 K is essentially due to the increase in the equilibrium constant. Therefore, if the combination of rate and equilibrium constants in eq 3 is to be represented by an Arrhenius form, the net activation energy must be near zero between 555 and 589 K. This would not explain the increase in the rate constant from 533 to 555 K. The temperature effect on the rate data
Table 111. Results of Regression Analysis eq 1
T. K
k x lo7 g-mol/@g of cat. atm)
c(residuals)*a x 1014
533 555 57 5 589
12.5 21.1 18.7 20.5
0.162 0.196 0.356 0.076
eq 3
R . E . . ~%
k X lo7 g-mol/(s g of cat.)
x(residualsY a
R.E..b %
16.5 23.5 26.0 15.5
6.22 14.14 12.62 13.80
0.210 0.138 0.263 0.054
20.5 19.0 24.6 13.7
Residual = rqXpu- rpred. R.E. = [ z$,Fexpu - rpWdl/rexpu)i] /N; rexptl = experimental rate; rpmd = rate predicted from regression equation; N = number of data points.
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983 643
could possibly be accounted for by retaining terms from the denominator of eq 2. They would have to decrease with temperature more sharply than the numerator and be comparable to unity in the 555-589 K temperature range but much larger than unity for temperatures less than 555 K. A t most three terms were retained due to the limited amount of data. The appropriate constants were determined by linear regression. For example, if the first three terms in the denominator of eq 2 are retained, the equation may be cast into the form
y=a
PC + bPc + c PH
where
a = K 1 K 2 / k owith k" = k3CT;b = K,/k"; and c = l/ko. The first requirement for a satisfactory model is that the regression analysis should yield positive constants. Furthermore, k" and K2 should increase with temperature and K1 should decrease with temperature. If these conditions are satisfied, the results of a linear regression analysis could be used as initial estimates to obtain improved values for the constants by nonlinear regression. Regression analyses were carried out by testing all possible combinations of two or three terms in the denominator of eq 2. The results displayed at least one negative constant over the temperature range of interest. We do not, however, feel that we have sufficient grounds to conclusively reject the suitability of the simplified models based on eq 2 as the constants become very sensitive to experimental error when the rate data are sparse. Again, due to the limited data, the calculated low relative error (in the order of 3 4 % if three terms in the denominator are retained) is not meaningful when the rate or equilibrium constants are negative. We can note here that with eq 1 which is based on the dissociative adsorption mechanism, the temperature effect on the rate cannot be explained if an Arrhenius type expression is assumed to represent the dissociative adsorption rate constant. In conclusion, at atmospheric pressure, over the temperature range of 533-589 K and under stoichiometric conditions with no diluent, there is evidence that a reaction mechanism where the rate-determining step is the 7 - c shift of the adsorbed cyclohexene is more suitable than a mechanism where the dissociative adsorption step is rate controlling. The ability of the simplest form of the rate expression based on the former mechanism (eq 3) to correlate the rate data is illustrated in Table 111. One can conclude that eq 3 is capable of representing the rate data adequately over a temperature range of 555-589 K with a constant value for the reaction velocity constant. Below 555 K inclusion of the adsorption terms will be necessary to correlate the rate data. Side Reactions The existence of side reactions in a thermochemical cycle presents potential problems in terms of catalyst deactivation and a decrease in the net thermal capacity of the cycle. Such technical problems may in most cases be resolved with appropriate intermediate purification steps and catalyst regeneration facilities. However, substantial economic penalties would be incurred in the form of both higher capital and operating costs which must be considered in the economic evaluation of any thermochemical cycle.
As the thermochemical cycles are closed systems, impurities in the reagents may be eliminated prior to the initial charging of the cycle. The question of extraneous species can then be limited to sources which are inherent in the chemistry of the system. With the cyclohexane system, the acid centers in the RD-150-C catalyst can provide a doorway through cyclohexene to a series of compounds based upon methyl cyclopentane and a family of n-hexane derivatives (Sinfelt, 1964). Cracking may also take place (Ritchie and Nixon, 1968) as well as an increase in the carbon number (Anderson, 1973). As indicated in the previous section, such observations have usually been made under conditions which are substantially different from those reported here. In the present case, there were in fact extraneous peaks present in the chromatographic data. The influence of the distribution of residence times in the recirculation reactor on the appearance of extraneous products was particularly evident in those instances where the reactor was bottled up in order to speed the approach to higher conversions. In these cases side products which were produced during the no-flow period of operation were gradually washed out of the reactor before the new steady state was reached. In reviewing these observations with respect to the intended applications, it is important to consider the distribution of residence times in the experimental reactor which would tend to enhance the impact of side reactions with large time constants. Such reactions would not necessarily play an important role in tubular converters which are being proposed for thermochemical cycles. At temperatures less than 672 K, two extra peaks appeared on the chromatogram, although not consistently. However, in all cases these extra products constituted less than 0.05% of the total area that was integrated. At 672 K a change took place. Two peaks were present but their percent of the total area increased to a maximum of 0.86% and their relative retention times changed in both magnitude and uniformity. Additional data at 700 K displayed three and four extra peaks on several occasions with further changes in magnitude and uniformity. The data beyond 589 K cannot be compared in every respect with the data at the lower temperatures as the relative retention times were, to some extent, a function of concentration (area percent) and in these runs the chromatographic column was subject to temperatures beyond its design limit. This caused a gradual deterioration of performance which necessitated individual calibrations for all of the kinetic data points. However, the increase in the area percent and the number of extraneous peaks are considered to be significant. The extra peaks in these data all appeared before benzene which appeared in all cases after the cyclohexane peak. Additional chromatographic evidence was obtained which indicated the presence of even lighter materials in the reaction products. A mass spectrometric analysis of a gas sample taken from the reactor operated at 589 K and 84% conversion is given in Table IV. The N2,02, Ar, and COz constituents are due to the sampling technique. The presence of butene and either butane or methylpentane in the sample led to a series of chromatographic and kinetic experiments specifically oriented to side product identification. The mass spectral technique could not distinguish between butane and methylpentane because methylpentane is fragmented during analysis and the corresponding peaks all appear in the cracking pattern of butane. Methylpentane happens to be the only C6with a cracking pattern very near to that of butane. Although there are subtle differences between the cracking patterns of methylpentane and butane, they
644
Ind. Eng. Chem. Process Des. Dev., Vol. 22, No. 4, 1983
Table IV. Mass Spectrometric Analysis of Gas Product (x = 0.84; T = 589 K) vol % constituent nitrogen 1.8 oxygen 0.10 argon 0.009 carbon dioxide 0.008 butene
cyclohexane benzene
toluene unknowna hydrogen a
0.16 0.23 0.26 0.0005 0.10 balance
An aliphatic hydrocarbon, possibly butane or
methylpentane.
would be indistinguishable in a mixture of low butane or methylpentane concentration. It was found that the reagent grade cyclohexane exhibited two extraneous peaks which appeared at 8 s and 11 s before the cyclohexane peak. Butane, 1-butene, and cis-trans-2-butene were individually added to cyclohexane and were all found to appear approximately 8 s before cyclohexane. It thus appears that butene and/or butenes are present in the cyclohexane feed stream. Methylpentane was also added to cyclohexane and was found to have a retention time 4 s before cyclohexane at low attenuation. Consequently, at normal attenuations, any methylpentane would be effectively masked by the cyclohexane peak on the chromatograph. The reagent grade cyclohexane was then introduced into the reactor and chromatograms were obtained at low attenuation over the range 533-589 K. The impurities which were present in the feed with relative retention times of -8 and -11 s were present in the reactor effluent. However, as the temperature increased the butane/ butene peak diminished and the area of the lighter impurity increased. Also, two lighter fractions progressively appeared as well as the heavier component. The heavier fraction was identified chromatographically as toluene. Apparently, butane and/or butenes entered the reador in the cyclohexane feed and were cracked or hydrogenated. The presence of butane as well as the hydrogenation possibility would explain the butane which was suspected in the mass spectrometric analysis. The cracking possiblity would explain the increase in lighter fractions as the temperature was increased as well as the appearance of toluene. The cracked products would encounter a plethora of benzene on the catalyst surface produced by the primary reaction. The cracking route would imply the area percents of the lighter fractions and of toluene would grow more rapidly than the decrease in the butane/butene peak. This, in fact, took place. Therefore, the observed impurities may be due in large part to the effects of feed impurities which must be substantially reduced in any further studies prior to the entry into the reactor system. It is pertinent to recall at this juncture that no loss in catalyst activity was found in this study and that the Engelhard RD-15OC catalyst contains acid centers which are known catalysts for other potential reactions. One could than conclude that irreversible side reactions are not significant over the range of experimental conditions reported here and that operation of thermochemical cycles under these conditions will not encounter substantial deteriorations in performance. There is, therefore, justification for building and testing a pilot scale loop to continuously cycle the cyclohexane, benzene, and hydrogen over the reforming and hydrogenation catalysts packed in single pass converters and investigating the
catalyst stability and reversibility of side products over an extended period of time. Acknowledgment The Union Carbide Corporation and Dr. J. M. Berty, currently affiiatd with University of Akron, are gratefully acknowledged for donating the internal recirculation reactor and Dr. R. Yarrington and Engelhard Industries for donating the catalyst and for valuable advice. The work was primarily supported by ERDA (E(ll-1)4031) and by Brookhaven National Laboratories (EC-77-C-02-4583). Nomenclature a = apparent particle sue defined as the ratio of the apparent catalyst volume to the external surface area C = molar concentration in gas phase C,, = surface concentration of cyclohexane CT = total catalytic site concentration De = effective diffusivity of cyclohexane in catalyst particle E = activation energy E = catalyst effectiveness k = reaction velocity constant k3 = velocity constant for the rate determining r u shift step k = velocity constant for dissociative adsorption step k8 = velocity constant defined as klCT K1...K, = equilibrium constants for the individual steps indicated in the reaction mechanism Kp = equilibrium constant C = catalytic site P = partial pressure PT = total pressure r = reaction rate per unit mass of catalyst R = ideal gas constant R = equivalent catalyst radius defined as 3a ?p= absolute temperature x = conversion of cyclohexane Greek Symbols AGR = free energy change for reaction A H R = enthalpy change for reaction ASR = entropy change for reaction fl = dimensionless parameter defined by (-AHR)E,C,/hT, y = dimensionless parameter defied by E / R T s h = thermal conductivity of the catalyst p = apparent density of catalyst particle Subscripts B = benzene C = cyclohexane eq = equilibrium value H = hydrogen s = property evaluated on the catalyst surface Superscripts 0 = standard thermodynamic conditions Registry No. Cyclohexane, 110-82-7. Literature Cited Anderson, J. R. A&. Catal. 1973, 23, 1. Berty, J. M.; Hambfii, J. 0.;Malone, T. R.; Ulbck, D. S. “Reactor for Vapor Phase Catalytic Studles”; 84th National AlChE Meeting, New Orleans, LA, Preprint 4ZE, 1969. Bridges. J. M.; Houghton, G. J. J . Am. Chem. Soc.1959, 8 1 , 1334. Butt, J. B. “Reaction Kinetlcs and Reactor Design”; Prentice HaH: Englewocd Cliffs, NJ, 1980 p 388. Cropley, J. B. Union CarbMe Corporation, South Charleston, WV, personal communication. 1981. Gerrnain. J. E. “Catalytic Conversion of Hydrocarbons”, Academlc Press: New York, 1969: p 108. a n d , J. L.; Baron, K.; SomorJai.G. A. J . Catal. 1975, 36. 305. Haensel, V.; Donaidson. G. R.; RW, F. J. “Thkd International Congress on Catalysis”; Amsterdam, 1984 pp 294-307. Johnson, J. R.; Rei&, J. J. “The Metal Hydride Development Program at Brookhaven Nationel Laboratory”: Proceedings of the DOE Chemical Energy Storage and Hydrogen Systems Contracts Review, Nov 1977, Jet Proputsion Laboratwy Pub. No 78-1 p 171, Feb 15. 1978. Mar, R. W.; Bramlette, T. 1.“Thermochemical Energy Storage Systems-A Review”, S a m Laboratories Energy Report, SAND 77-8051, prepared under DOE Contract AT (29-1) 789, Livermote, CA, Feb 1978. Rltchle, A. W.: Nlxon, A. E. I d . Eng. Chem. prod. Res. Dev. 1968, 7 , 209.
Ind. Eng. Chem. Process Des. Dev. 1983, 22, 645-653 Ruiz-Vizcaya, M. E.; Novaro, 0.;Ferreira, J. M.; Gomez, R. J . Catal. 1978, 51, 108. Sinfelt, J. H. A&. Chem. Eng. 1964, 5 , 37. SRtig, M. "Catalysts and Catalytic Processes": Chemical Process Revlews, Noyes Development Corp.: New Jersey, 1967: p 57. Smith. R. B. Chem. €no. Roo. 1959. 55f61. 76. Smith; R. L.; Prater, C. 6. Chim. €ng.'R&.'Symp. Ser. 1987, 73(63), 105. Tetenyi, P. Acta Chim. Budepest. 1974, 82, 459. Thomas, C. L. "Catalytic Processes and Proven Catalysts"; Academic Press: New York, 1 9 7 0 p 399.
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Received for review September 23, 1981 Revised manuscript received June 3, 1982 Accepted March 30, 1983
Application of a Catalyst Deactivation Model for Hydrotreating Sohrent Refined Coal Feedstocks Ramakrlshna V. Nalltham, A. Ray Tarrer,' James A. Guln, and Christine W. Curtls Auburn Coal Conversion Laboratory. Chemical Engineering Department, Auburn University, Auburn, Alabama 36849
A simple kinetic model, including a first-order catalyst deactivation rate, is applied to upgrading of cualderived feedstocks prepared from two solvent refined coal fractions. A catalyst deacttvation mechanism is proposed which involves the adsorption and surface reaction of coke precursors on catalytic active sites. The effect of feedstock composition, temperature, and pressure on kinetic parameters, and in particular, the catalyst deactivation rate,
is determined.
Several processes have been developed over the past few decades for conversion of coal to clean solid and liquid fuels. Hydrotreatment involving catalytic hydrogenation and hydrocracking of the heavy coal liquids using a hydrogen rich gas is generally utilized to upgrade the initial coal liquids. The activity of the catalysts used in the hydrotreating step generally diminishes with time and hence, the product quality as measured, for example, by degree of asphaltene conversion to oil, or heteroatom removal, deteriorates. Over a limited period, this decline in catalytic activity can be offset by operating the unit at higher severity; however, eventually the spent catalyst must be replaced or regenerated. The cost of commercial hydrotreating catalysts has escalated significantly over the past few years and the economic feasibility of hydrotreating processes is affected by the life of the catalyst. In addition to this, a knowledge of the rate of catalyst deactivation is essential for the design of the hydrotreater and for estimating the process severity-time relationship required to achieve a constant conversion of reactants to products. The objective of the present work is to examine the initial rapid deactivation of a hydrotreating catalyst in upgrading coal liquids in light of the kinetics of the upgrading process. A mechanism for the deactivation is proposed and rate parameters in the deactivation rate model are related to feed properties and processing conditions. Deactivation of hydrotreating catalysts in the presence of coal liquids is a very complex phenomenon due to a wide variety of components in the feedstock. A hydrotreating catalyst can suffer deactivation by many mechanisms including fouling, poisoning, sintering, and loss of sulfur from the catalyst. In addition, pore-plugging can reduce the intraparticle diffusion of reactants into the catalyst pores and hence reduce the rate of reaction. The exact mechanism causing the deactivation may vary depending on the catalyst age. Sie (1980) observed three distinct stages of 0796-4305/83/7 l22-0645$07.50/0
deactivation in a study of hydrodesulfurization of a residual feedstock. A rapid decline in activity was observed during the initial and final stages and a gradual decline of activity was observed during the intermediate stage. The rapid initial activity decline was attributed to carbon deposition, while the slow decline was attributed to poisoning by metals such as vanadium and nickel. Stiegel et al. (1982) studied catalyst deactivation during coal liquefaction. The deposition of coke on catalyst is rapid within the first few hours of processing and contributed to the initial decline in activity. The results indicate that metal deposition is the major cause of longterm catalyst deactivation. Ocampo et al. (1978) reported rapid and severe decline in the hydrogenation activity of hydrotreating catalysts in batch coal liquefaction experiments within the first few hours of processing. Mitchell (1980) compared the hydrogenation and hydrocracking activities of fresh and deactivated catalysts and observed that both declined significantly after deactivation and that the hydrocracking activity declined more rapidly than the hydrogenation activity. Furimsky (1978) observed significant quantities of coke formed on catalysts during hydrotreatment of bitumen and heavy gas oil in the initial period of processing. The mechanism of coke formation on hydrotreating catalysts is complex and not well understood. Beuther et al. (1980) proposed that the mechanism of coking was similar to the phenomenon of mesophase formation in carbonization of aromatic liquids except that the smaller mesophase crystals were converted rapidly to coke before they coalesced to form larger crystals. In its simplest form, coke formation on the catalyst can be viewed to result from the adsorption of coke precursors on the active sites with subsequent participation in degradation reactions such as condensation, dehydrogenation, and polymerization, to form high molecular weight species (Corella and Asua, 1982). This mechanism is similar to that presented by 0
1983 American Chemical Society
