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Ind. Eng. Chem. Process Des. Dev. 1986, 2 5 , 278-284
Desulfurization of Dibenzothiophene by in Situ Hydrogen Generation through a Water Gas Shift Reaction Bruce D. Hookt and Aydln Akgerman’ Kinetics, Catalysis, and Reaction Engineering laboratoty, Chemical Engineering Department, Texas A&M University, College Station, Texas 77843
Global depletion of sweet crude necessitates greater use of high sulfur content hydrocarbon resources. Because of this and the prohibitive cost of producing pure hydrogen, hydrodesulfurization is becoming a major cost factor in petroleum and coal liquid refining. This study indicates that hydrodesulfurizationcan be achieved by generating the hydrogen necessary for the reaction in situ with the water gas shift reaction. For this work, a sulfided commerical Co0-Mo03/y-AI,03 catalyst is used in a trickle-bed reactor at 58-69 atm of pressure and 595-648 K temperature. Dibenzothiophene is chosen for this work since it is representative of the heterocyclic sulfur molecules present in oil which are rather difficult to desulfurize. Significant desulfurization of dibenzothiophene by in situ hydrogen generation is observed. The pseudo-first-order rate constants calculated appear to be an order of magnitude higher than those reported in the literature for pure hydrogen feed. Zero-, first-, and second-order rate expressions are tested with the data analysis program developed previously, and activation energies and rate constants are calculated for each case. The three expressions have comparable validity based on the sum of the squares of deviations between predicted and experimental conversions. The activation energies calculated agree well with literature values published.
Hydrodesulfurization (HDS) of crude oils and coal liquids has gained increasing importance over the last several years as low sulfur content hydrocarbon resources have been depleted and become more expensive. Researchers have concentrated on developing catalysts and investigating the kinetics and mechanisms of the HDS reactions. Thiophenic sulfur is the hardest to desulfurize (Gates and Schuit, 1973). Heterocyclic thiophenic compounds are among the least reactive of the sulfur-laden compounds found in petroleum residua, tar sands bitumen, and coal-derived oils. Recent studies have shown that dibenzothiophene (DBT) is one of the most difficult to desulfurize polycyclic thiophenes that has been studied (Nag et al., 1979; Singhal et al., 1981a). For early researchers, hydrogenation of one benzo ring seemed necessary for the HDS of dibenzothiophene (Cawley, 1951). The product distributions of more recent works have depended heavily on the temperature and pressure of the system studied, the flow conditions (i.e., whether the reaction occurred in the gas or liquid phase), and the type and activity of the catalyst used. Studies done a t atmospheric pressure resulted in biphenyl (BP) being the predominant reaction product for DBT desulfurization (Bartsch and Tanielian, 1974; Kilanowski et al., 1978). Studies done at higher pressures yield cyclohexylbenzene (CHB) and bicyclohexyl (BCH) as additional significant reaction products. Hoog (1950) obtained product distributions containing as much as 65% selectivity toward CHB and BCH at 648 K and 50 atm. Rollman (1977) found that product selectivity a t low temperature (617 K) leaned almost exclusively toward biphenyl. A t high temperature (672 K), however, CHB production had increased as high as 25% of the converted DBT. These results were generally confirmed by others (Houalla et al., 1978; Nag et al., 1979; Geneste et al., 1980; Singhal et al., 1981a,b). The two most current mechanisms for dibenzothiophene HDS have been proposed by Houalla et al. (1978) and by Singhal et al. (1981a,b). The Houalla et al. mechanism At present in the Chemical Engineering Department, Rensselaer Polytechnic Institute, Troy, NY 12181. 0196-4305/86/1125-0278$01.50/0
indicates a combined parallel-sequential model in which biphenyl and hydrogenated dibenzothiophenes are the primary products. CHB is formed, according to this model, from either of the primary products and may hydrogenate further to BCH. However, very little BCH was detected by Houalla et al. They have presented rate constants that are pseudo-first-order with respect to the organic reagent and determined at 573 K, 102 atm, with a sulfided CoOMOO,/ y-Al,O, catalyst. The mechanism for DBT hydrodesulfurization proposed by Singhal et al. (1981a) is more complex. Their mechanism proposes initial hydrogenation of one of the double bonds of the thiophene ring. This is followed either by further hydrogenation of the benzo ring to give tetrahydrodibenzothiophene (THDBT) or hexahydrodibenzothiophene (HHDBT) or by the scission of a C-S bond to form a mercaptobiphenyl. The mercaptyl group is quickly extracted to form biphenyl. THDBT and HHDBT are desulfurized by breaking the C-S bond next to the hydrogenated ring to form a mercaptan intermediate which is desulfurized quickly to CHB. Kinetic studies of dibenzothiophene HDS by Singhal et al. (1981b) also indicate that a finite portion of the CHB found in the HDS products comes from biphenyl hydrogenation. No BCH, however, was found by Singhal et al., probably because the hydrogen concentration was not as high as that of Houalla et al. (1978) and others (Hoog, 1950; Landa and Mrnkova, 1966). Singhal et al. (1981a,b) studied DBT desulfurization at 558-623 K, 31 atm, in the gas phase over a standard Co0-MoO3/y-Al2O3catalyst. Both of these mechanisms are consistent with the generalized mechanism for HDS proposed by Kwart et al. (1980). Their mechanism suggests that the C-C bond next to the sulfur atom in the thiophene ring hydrogenates first, followed by either C-S bond cleavage or by further hydrogenation and then C-S bond cleavage. Cobalt-molybdenum catalysts have been the workhorse of HDS processes for years. De Beer et al. (1976) have shown that nickel- or cobalt-promotedmolybdena-alumina catalysts are more active than tungsten sulfide or molybdenum sulfide or Mo-A1 alone. Although presulfiding may not be essential for industrial HDS catalysts, since the oxidic form is transformed to the sulfide form during the 0 1985 American
Chemical Society
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Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 1, 1986
reaction (McKinley, 1957), the steady-state activity of the catalyst may not be reached for some time if presulfiding is not performed (Gissy et al., 1980a; Daly, 1978). Gissy et al. (1980a) found that benzothiophene HDS did not reach a steady-state rate until nearly 20 h on stream had passed when the catalyst was initially in its oxidic state. Prereducing the catalyst with hydrogen before starting the HDS reaction did not improve the activity much over the oxidic form. Presulfiding, however, significantly increased the initial activity of the catalyst and shortened the unsteady-state time of the catalyst activity by more than half. Gissy et al. (1980b) reports that presulfding with a HzS/Hz mixture was the best technique for activating a COOMo03/r-Alz03catalyst for HDS. An HDS reaction with in situ hydrogen generation has already been achieved by Kumar et al. (1984). As much as 60% desulfurization of benzothiophene with a commercial Ni0-Mo03/r-A1,0, catalyst a t 68 atm and 583-633 K was obtained. Other work by Takemura et al. (1983) showed that hydrogenation of nitrobenzene could be accomplished with in situ hydrogen generation from the water gas shift reaction. However, these reactions are known to be faster and easier to accomplish than DBT hydrodesulfurization. The trickle-bed reactor is the most prevalent reactor used in hydrodesulfurization and most other hydrotreating reactions. Most of the research on trickle-bed reactors has been performed near ambient conditions. At these low temperatures and pressures, the liquid phase can generally be assumed to be nonvolatile. However, since most hydrotreating processes operate a t conditions much more severe than that-up to 700 K and 200 atm-the reactants will almost surely flash to some extent (50% flashing is not uncommon in industrial hydroprocessors). The presence of reactant in both the gas and liquid phases significantly complicates the reaction scheme, especially if the catalyst is partially wetted. Herskowitz and Smith (1983) assert that reactions for which the rate is controlled by a nonvolatile reactant will undergo a reduction in the rate if partial wetting occurs. If the limiting reactant is in both phases due to vaporization of the liquid, however, and mass-transfer effects are significant, then partial wetting results in an increased reaction rate. Sedricks and Kenney (1973) found that for the reduction of crotonaldehyde, the reaction occurring on the Udryn parts of the catalyst dominated the overall reaction rate, even when the wetting efficiency was large. Another consequence of partial vaporization may be observed in the analysis of data taken from trickle-bed reactors. Collins et al. (1984) showed that some of the data published for high-temperature/ high-pressure trickle-bed operation actually flashed completely to the vapor phase. Hook et al. (1984) have presented a technique for extracting intrinsic kinetics from nonisothermal trickle-bed reactors operating a t high temperatures and pressures. The Hook et al. technique uses the SoaveRedlich-Kwong equation of state to predict the distribution of the components between the phases a t collocation points throughout the reactor. A Gaussian quadrature method is used to integrate the plug flow design equation using axial temperature profiles obtained during experimentation. Hook and Akgerman (1984) have shown that this technique can be used to accurately estimate the Arrhenius constant and activation energy for the benzothiophene HDS data published by Kumar et al. (1984).
Experimental Section A schematic of the apparatus used in this study is presented in Figure 1. The reactor used was a 63.5 cm
279
E
>
A . Water Feed
B . Hydrocarbon Feed
I
C Pressure G a u g e 0 . A i r To M i x S a l t B a t h
S . Gas Sampling Knockout T . Separator Used U n t i l Steady State
U.
0 8 8 Sampling P o r t V . Bubble Meter W . Water Scrubber x . vent
X
Figure 1.
long, 1.91 cm 0.d. stainless steel seamless tube placed vertically in a 45.72 cm deep (10.23 cm i.d.) bath filled with a molten eutectic salt. The reactor tube had an inside diameter of 1.575 cm and was packed with 0.64 cm of glass wool at the bottom outlet, 8.73 cm of inert ceramic material on top of the glass wool, and 33.81 cm of catalyst and then was filled to the top with inert ceramic. The dibenzothiophene (DBT) was dissolved in decalin and fed to the reactor a t feed rates ranging from 4.48 to 8.49 mL/min by using a Pulsafeeder Microflo metering pump. The concentrations of DBT in decalin used were 3.36-3.62 wt %. Distilled and deionized water was fed to the reactor by a Milton Roy Minipump at rates of 0.243-0.426 mL/min. Carbon monoxide (CO) flow was monitored and controlled by a Model 5850 Brooks Instrument Co. mass flow meter. The CO flow rates used in this study were 400-620 std cm3/min. The hydrocarbon and water feeds were heated simultaneously with a tube wrapped in heating tape prior to entering the reactor a t the top. The gas feed was passed through a coil submerged in the molten salt bath and then introduced to the hydrocarbon and water feed upstream of the reactor entrance. Both the gas and the liquid phases exit the reactor at the bottom and passed through a water-jacketed cooling section and the back-pressure regulator to the collection systems. During the time when samples were collected for product distribution analysis, the product stream was diverted to water-jacketed-phase separators which collected the liquid phase while allowing the gas-phase components (now at atmospheric pressure and room temperature) to pass through. A gas sampling port at the outlet of the third-phase separator allowed gas and liquid samples to be obtained simultaneously. During the time it took to reach steady state, when samples were not being collected, the liquid products accumulated in a large separator, while the gas products passed through to a “knockout pot” which sparged the gas through an aqueous sodium hydroxide solution (pH >8.0) to remove H2Sbefore venting. Materials. Dibenzothiophene and cyclohexylbenzene were obtained from Alfa Products. Bicyclohexyl and biphenyl were acquired from Aldrich Chemical Co., and decalin was bought from Pfaltz & Baur, Inc. All products were used as received except the decalin, which was filtered to remove particulate matter. Chromatographic analysis of the decalin showed that it contained some tetralin and naphthalene. No attempt was made to remove these compounds from the decalin since they did not affect the HDS reaction product analysis and their relative concentration did not change after the reaction. This indicates that no hydrogen was consumed by naphthalene or tetralin
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Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 1, 1986
hydrogenation. Likewise, it indicates that decalin did not dehydrogenate to act as a hydrogen donor. Catalyst. The catalyst bed consisted of 33.42 g of Co0-MoO3/~-AlZO3 catalyst (0.16 cm of Aero HDS-2A from American Cyanamid Co.) The catalyst composition was Moo3 15.4%, COO3.2%, SiOp0.1%, and A120380.9%, as received with a BET area of 310 m2/g and pore volume of 75 cm3/g. The average length of a randomly chosen sample of the catalyst (0.160 cm diameter extrudate) was found to be about 0.5 cm; the inert ceramic material was ground ceramic beads sieved to yield a particle size range of 0.084-0.119 cm. The catalyst was pretreated before use. First, air was passed over the catalyst bed at 660 K for 4 h. Second, a 10% HzS-90% H, by volume gas mixture was passed over the catalyst for 8 h a t atmospheric pressure and 630 K at a rate of 400 std cm3/min (sccm) in order to presulfide the active sites of the catalyst. Following the presulfiding, the exit gas was compared with the feed presulfiding gas to ensure “breakthrough”. Then hydrogen was passed over the bed for 2 h a t 400 sccm to remove excess HzS. A t no time after pretreatment was the catalyst bed exposed to air or oxygen. Procedure. Typical experiments were done in groups of four flow rates at one salt bath temperature. A t the beginning of a run of four experiments, the reactor was pressurized with nitrogen and flooded with decalin after it reached the salt bath temperature. This was done to ensure that the pores were completely filled with liquid, and the catalyst surface started out completely wetted. After 15 min or more, the outlet valves were reopened, CO flow was started, the water pump was turned on, and the hydrocarbon pump was switched from decalin to the DBT in decalin solution. The reactor was allowed to run for an hour or more to allow it to come to steady state. One hour of operation was approximately 17 residence times a t a liquid flow rate of 8.5 mL/min. Steady state was known to exist when the temperature profile stabilized and by the number of residence times that had passed since the system had been disturbed. Steady-state samples of the liquid and gas products were obtained by diverting the product stream to the steady-state collection/separation system. During the time that liquid product was collected (3-5 min), the temperature profile in the reactor was recorded and a 10-20-mL gas sample was taken a t atmospheric pressure. Generally, three steady-state liquid samples were obtained during the course of an experiment. Also, after the fourth flow-rate experiment had been completed, the hydrocarbon feed rate was set at the pump setting used for the first experiment of that run. The reactor was allowed to come to steady state, and one sample was taken to ensure reproducibility. In order to bring the catalyst to steady-state activity, on and off runs were performed for about a week, a total of 40+ h. No change in catalyst activity was noted thereafter. When the reactor was not running, the catalyst bed was kept under an atmosphere of inert nitrogen. Product Analysis. Both gas and liquid samples were analyzed by gas chromatography. The gas samples were analyzed on a Carle gas chromatograph which utilized a complex series of columns in which the flow and temperature were controlled by the Hewlett-Packard 3385 automation system. This system was described in detail by Phillip et al. (1979). A Varian Vista 44 gas chromatographic system was used to analyze the liquid samples for their product distributions. The column used in the Varian chromatograph was a 25 m long, 0.33 mm i.d., QC3-BP10 (dimethyl phenyl
cyano OV-1701) wide-bore vitreous silica capillary column from Scientific Glass Engineering, Inc. Results a n d Discussion The primary objectives of this work were (1) to show that the hydrodesulfurization of dibenzothiophene (DBT) could be accomplished by generating the hydrogen necessary for the reaction in situ through the water gas shift reaction and (2) to study the kinetics of the HDS reaction in terms of total sulfur removal and to determine the best rate expression for the system studied. A secondary objective was to study the product distribution and compare it to published data. The technique developed by Hook et al. (1984) which takes into account the partial vaporization of the liquid phase was used in data analysis. The overall rate of sulfur removal as a single reaction was studied, and a simple model was used which did not consider possible parallel or sequential paths of desulfurization. This technique has been employed successfully to fit the data of Kumar et al. (1984) as presented in Hook and Akgerman (1984). The technique assumes (1)that all the reaction occurs in the liquid phase, (2) that the catalyst is completely wetted, (3) that the liquid phase is in plug flow, (4) that all mass- and heat-transfer effects are negligible, i.e., that the vapor and liquid phases are in physical equilibrium, and (5) that radial concentration and temperature gradients are negligible. The assumption of complete wetting is generally valid for porous catalysts when a liquid phase is present to any significant extent. The liquid-phase wetting efficiency is based on the surface area of the catalyst in contact with the liquid phase. The assumption is reasonable since most of the surface area of the catalyst used is internal, and the internal wetting efficiency is unity or nearly unity over a wide range of liquid flow rates (Mills and Dudukovic, 1981). The Lee and Smith (1982) criterion also was used and did not indicate incomplete wetting for any of the cases tested. For all cases, the Thiele modulus was less than 0.2, indicating no transport effects. A list of the flow rates, total pressures, and average temperatures studied is presented in Table I. The reaction conversions given are based on total sulfur removal and on conversion of dibenzothiophene. A graph of sulfur removal as as function of inverse sulfur weight hourly space velocity (SWHSV) is shown in Figure 2. The scatter of the data in Figure 2 is due to the existence of significant temperature profiles in the catalyst bed for each run. These profiles may be found elsewhere (Hook, 1984). Because of these temperature profiles, this graph is not as representative as one which displayed data from an isothermal reactor, but it is presented in order to give some idea of the effect of flow rate on conversion. Unfortunately, available reactor dimensions limited the size of the catalyst bed which could be used. This in turn limited the extent of reaction possible. The reaction equation for a plug flow reactor may be written as -d n, - = r, = k , f ( c j ) dW Conversion, x,, may be expressed as x, =
(ni - n,)/G
(2)
and the rate constant in the usual Arrhenius form
k , = A exp(-E/Rn
(3)
The Arrhenius form of the rate constant does not converge
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 1, 1986
281
Table I. Experimental Flow Rates and Conversionsn run 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21
Ptotal
T.¶"
62.2 62.2 62.2 69 69 62.2 62.2 62.2 62.2 62.2 61.5 62.5 62.2 63.2 62.2 62.2 62.2 62.2 58.5 58.5 58.5
630.2 628.7 635.9 617.2 614.3 614.7 614.1 623.3 622.3 623.2 622.7 643.1 639.9 639.7 638.6 629.8 633.4 614.2 621.7 622.4 622.1
102na 1.351 1.351 1.492 2.354 2.354 2.354 2.354 1.496 1.539 1.513 1.520 1.533 1.535 1.540 1.506 1.543 1.537 2.369 1.463 1.467 1.457
1 0 ~ ~ ; 102n&c ~ ~ 2.940 5.090 4.098 2.370 3.141 1.810 3.036 0.792 3.504 0.883 4.096 1.068 1.260 4.828 2.801 0.784 3.577 1.000 4.076 1.140 1.350 4.838 3.536 0.990 1.150 4.106 4.668 1.310 5.131 1.440 3.633 1.010 4.080 1.130 3.350 0.940 3.698 1.030 4.028 1.120 4.651 1.290
102nlo 2.759 2.759 2.759 2.802 2.802 2.802 2.802 1.829 1.829 1.825 1.825 1.820 1.820 1.820 1.820 1.816 1.820 2.757 1.820 1.820 1.820
~DBT
Xl
0.0616 0.0368 0.0934 0.1321 0.0886 0.0756 0.0581 0.1640 0.1301 0.1189 0.0917 0.3133 0.2660 0.2327 0.1916 0.2582 0.2540 0.1095 0.1909 0.1541 0.1293
0.0812 0.0475 0.1247 0.1564 0.1080 0.0955 0.0709 0.1918 0.1541 0.1452 0.1119 0.3453 0.2956 0.2604 0.2161 0.2836 0.2791 0.1278 0.2079 0.1687 0.1429
XWGS
0.9039 0.9856 0.9798 0.8791 0.8598 0.8231 0.7707 0.8920 0.8578 0.9273 0.8650 0.8909 0.8734 0.8843 0.8737 0.8634 0.8731 0.7680 0.8864 0.8513 0.8702
"The units for the molar flow rates were gmol/min. For temperature, the units were K. Conversions were dimensionless. Pressure units are atm.
of the rate parameters. Therefore separating variables and integrating yields 0.354
0.301
Gaussian quadrature was used to evaluate the integral on the left-hand side and also to change the right-hand side into a summation from which the rate parameters could be evaluated numerically. Thus
0.25-
z
'C
0 v)
0.20
5
z
0
0 0.15
0.10
0.05
0.00
5
7
9
11
13
15
17
19
21
23
l . / S U L P H U R W H S V lHRl
Figure 2.
very well numerically due to the correlation between A and E. Therefore, a reparameterization is necessary.
Equation 1 may now be expressed as
For small temperature fluctuations, f(Cj) may be evaluated through the reactor at an average temperature. However, the rate constant is much more sensitive to the temperature profile in the reactor, and so the axial temperature profile measured in the reactor is included in the analysis
where npl, np2= the number of qusdrature points used in the evaluation of the integral and Cj(xJ, T(WJ = variables evaluated a t the quadrature points. For each quadrature point, a flash calculation was performed on the reaction system. First a material balance was performed on the system based on the sulfur conversion at the quadrature point. Next the flash routine developed by Collins (1983) calculated the distribution of components between the gas and liquid phases. The flash calculation also calculated the liquid density. From this information, the liquid-phase concentrations for all components were calculated (Hook et al., 1984). All the calculations done for the total sulfur removal analysis assume the water gas shift reaction to be in dynamic equilibrium. This assumption was valid because the gas-phase product analysis generally showed equilibrium conversion for the water gas shift reaction (within experimental error) and also because of the large increase (about 10-15 "C)in temperature at the beginning of the catalyst bed which leveled off after a few centimeters of packing. The highly exothermic water gas shift reaction would be likely to generate such a temperature profile if most of the reaction conversion occurred a t the beginning of the reactor. Three rate expressions were chosen to try and model the reaction in terms of total sulfur removal: zero order first order second order
rs
rs
rs
= ks0
=
ksOCbBT
=
ksOCLCbBT
282
lnd. Eng. Chem. Process Des. Dev., Vol. 25, No. 1, 1986
Table 11. Rate Parameter Estimates" reaction order k , a t To = 627 K zero 5.159 X lo4 f 0.317 X 10first second
SSR =
E , kcaligmol
3.660 X lo-' f 0.399 X lo-' 3.037 X 10' f 0.307 X 10' X(Xexpt1-
31.2 f 4.6 36.9 7.9 36.7 f 7.3
*
A 3.96 X lo5 gmol/(g of catalyst min) 2.64 X 10" cm3/(g of catalyst min) 1.92 X 1015cm6/(gmol g of catalyst min)
SSR 0.0300 0.1059 0.1098
Xpred)'
'The error bands indicated above are the asymptotic standard error.
Table 111. Predicted Conversions for Overall Sulfur Removal, x, run 1 2 3 4 5 6 " I
8 9 10 11 12 13 14 15 16 17 18 19 20 21
exptl 0.0616 0.0368 0.0934 0.1321 0.0886 0.0756 0.0581 0.1640 0.1301 0.1189 0.0917 0.3133 0.2660 0.2327 0.1916 0.2582 0.2540 0.1095 0.1909 0.1541 0.1293
zero 0.0661 0.0767 0.1331 0.1452 0.1153 0.0961 0.0807 0.1881 0.1405 0.1285 0.1059 0.3222 0.2455 0.2150 0.1882 0.1887 0.1946 0.1500 0.1071 0.1500 0.1343
order first 0.0849 0.1332 0.2921 0.1434 0.0951 0.0809 0.0477 0.1724 0.0954 0.0709 0.0786 0.3006 0.2141 0.1730 0.1433 0.1567 0.1607 0.1167 0.1076 0.1416 0.1058
second 0.1843 0.1446 0.3028 0.2137 0.1476 0.1127 0.1022 0.2277 0.1579 0.1850 0.1126 0.3779 0.2737 0.2369 0.2104 0.2132 0.2137 0.1619 0.1490 0.1869 0.1631
The rate constant, ko,activation energy, E , and Arrhenius constant, A , calculated as outlined above for each rate expression are given in Table 11. The details of the derivation of the function f(CJ are given elsewhere (Hook et al., 1984). Each rate expression tried and the activation energy and Arrhenius constant associated with it were tested in the trickle-bed model developed by Collins (19831, which has been modified slightly to improve its convergence characteristics. Comparisons of predicted and experimental values for conversion for the rate expressions tested are presented in Table 111. The sum of the squares of the deviation (SSR) in Table I1 show that there is little difference between the prediction errors between the models. The activation energies for all expressions are within range of each other, however, and this indicates that the firstand second-order rate expressions are probably not as inaccurate as the SSR indicates. Rate expressions for sulfur removal published in the literature have generally assumed a first-order expression (Bartsch and Tanielian, 1974; Kilanowsky et al., 1978; Houalla et al., 1978; Singhal et al., 1981a,b). However, most of the rate data in the literature have been obtained with one phase (usually a gas phase) in contact with the catalyst, or a nonvolatile liquid-phase assumption has been made (Scamangas et al., 1982; Qader et al., 1968). At the temperature and pressure conditions necessary to desulfurize dibenzothiophene, a nonvolatile liquid assumption clearly is not valid (Collins, 1983; Hook et al., 1984). One exception to assuming first-order kinetics is the work by Geneste et al. (1980) who found that a zero-order fit satisfactorily accounted for their DBT desulfurization rate data. Geneste et al. used a batch stirred reactor and although the amount of reactants charged to the reactor was not published, the data indicate the presence of two phases. Although Singhal et al. (1981b), working with a gas-phase
fixed-bed reactor, fit their data with a pseudo-first-order rate expression, they indicate a linear relationship between conversion and inverse space velocity for conversions less than 55%. Likewise, Houalla et al. (1978) fit their data to a linear conversion vs. contact time relationship for conversions greater than 60%. It is not surprising therefore that in our experiments, the zero-order rate expression also gave a slightly better fit. A zero-order rate expression implies that the surface concentration of the reactants is essentially constant throughout the reactor. Since the reaction rate is slow compared to the rate of diffusion, the surface concentrations of the reactants are equivalent to the bulk liquid concentrations of the reactants. A t low conversions, the surface concentrations do not change appreciably during the course of the reaction, and the change in conversion with space velocity is linear. In this work, that possibility is enhanced by the presence of both vapor and liquid phases in the reactor. At the temperatures and pressures studied, a significant fraction of the liquid-phase components will vaporize in the preheating section of the reactor. As a result of this partial vaporization, any DBT which flashes to the vapor phase at the beginning of the reactor will be available to replenish the liquid-phase concentrations further down the reactor. Because of this masking effect caused by partial vaporization and vapor-liquid equilibrium, and the statistically comparable validity of all three rate expressions tested, the zero-order rate expression should not be construed as indicating the mechanism of the reaction. It only shows the effect that partial vaporization can have on the global observed kinetics and further indicates the need for including VLE calculations in the data analysis techniques used for trickle-bed reactors. Unfortunately, because of limitations in size of the reactor, we could not achieve higher conversions. However, this does not limit the comparison of our results with those reported in the literature. Activation energies for DBT hydrodesulfurization presented in the literature range from 28 to 39 kcal/gmol. Geneste et al. (1980) found an activation energy of 28 kcal/gmol for their zero-order rate expression for DBT hydrodesulfurization using a sulfided Co0-Mo03/yA1,03 catalyst used industrially. Rollman (1977) reported an activation energy of 36 kcal/gmol for a rate expression which was second-order (first-order in both hydrogen and DBT) using the same catalyst as in this work (HDS-2 from American Cyanamid). Both of these activation energies have been for overall sulfur removal rate expressions. Singhal et al. (1981b) propose an activation energy of 39.1 kcal/gmol for the slow step of their kinetic model. This slow step is presumably the formation of a hydrogenated dibenzothiophene. Their rate constant however is based on Langmuir-Hinshelwood kinetics, although it is firstorder in both hydrogen and DBT. Singhal et al. also used a sulfided Co0-Mo03/yA1,03 catalyst. The zero-order activation energy found in this work (31.23 kcal/gmol) compares well with that obtained by Geneste et al. (1980). Likewise, the first- and second-order rate expression activation energies (36.9 and 36.7 kcal/
Ind. Eng. Chem. Process Des. Dev., Vol. 25, No. 1, 1986 283 Table IV. Comparison of Literature Rate Constants with Values Calculated from This Work pseudo-first-order rate const source Houalla et al., 1978 Rollman e t al., 1977 Singhal et al., 1981b Lee and Ollis, 1984
T, K 573
lit 2.8 x 10-5
617
7.5. 6.3
this work units 2.1 X m3/(kg of catalyst s) 110 h-’
625 0.1
2.0
598 0.0188 = k
~
gmol/(g of catalyst h) gmol/g of catalyst
~ 0.524 s
gmol, respectively) agree well with those published by Rollman (1977) and Singhal et al. (1981b). One possible explanation for the zero-order “best fit” rate expression and the lower activation energy associated with it, both in this work and in that of Geneste et al. (1980), may be the presence of two phases in the reaction system. Vapor-liquid equilibrium effects may be masking the true kinetic models for the reaction as explained above. Table IV compares the rate constants obtained in this work by employing the first-order fit with those available from the literature. It was not possible to make any comparison to zero-order rate constants cited by Geneste et al. (1980) since they did not specify the amount of initial charge and the amount of catalyst in their batch system. In order to have meaningful comparisons, the predicted rate constants evaluated with the rate parameters from this work have been converted to the same units used in the literature. As can be seen from the table, the rate constants from this work, i.e., HDS by in situ hydrogen generation, are an order of magnitude greater than those reported for HDS with pure hydrogen feed. The work reported in the literature, however, was done at or below the lower end of the temperature range of this work. Therefore, we believe that the comparison of the rate constants is not conclusive. The definite trend of higher rate constants, on the other hand, suggests that HDS by in situ hydrogen generation through the water gas shift reaction has a greater rate constant because the hydrogen formed a t the surface is nascent hydrogen and is more active. Most researchers acknowledge that for HDS reactions, the rate expression is also a function of the hydrogen concentration, which is first-order in most cases. However, most of them feed large excesses of hydrogen into the
reactor which makes detection of this dependence very difficult if not impossible’ (Weisser and Landa, 1973; Rollman, 1977; Houalla et al., 1978; Singhal et al., 1981b; Scamangas et al., 1982). This allows them to lump the hydrogen concentration into the rate constant and analyze their data with a pseudo-first-order rate expression. Including the hydrogen concentration term in the rate constant causes the activation energy to absorb whatever temperature dependence the hydrogen concentration term has. This term becomes especially significant in systems where a liquid phase is present. Observed products of the reaction of dibenzothiophene and hydrogen were H2S, biphenyl cyclohexylbenzene (CHB), tetrahydrodibenzothiophene (THDBT), and trace amounts of bicyclohexyl (BCH). the presence of these compounds was identified first by gas chromatography and later confirmed by mass spectroscopy. The product distributions observed in this work agree with those obtained by Singhal et al. (1981a)b)and Houalla et al. (1978). Biphenyl was the major product, with lesser amounts of CHB being produced. THDBT was found in all the samples tested. The ratio of THDBT to biphenyl varied as a function of space velocity and temperature. The product distributions from each run are listed in Table V. It is interesting to note that for all our runs, the ratio of biphenyl to CHB is about 9. This is the same result obtained by Singhal et al. (1981a,b). Conclusions The results presented above indicate that desulfurization of DBT by in situ hydrogen generation through the water gas shift reaction proceeds a t a similar rate as desulfurization by hydrogen only. In addition, the product distributions obtained agree well with those published by others who fed hydrogen only. Since the activation energies obtained in this work are in agreement with those published in the literature, it may be concluded that for HDS reactions, generating hydrogen in situ is probably as efficient and effective as feeding hydrogen to the reactor. The data published by Kumar et al. (1984) for benzothiophene HDS backs up this conclusion. In addition, two items were presented which merit further study. The first was the masking effect of partial vaporization followed by downstream replenishment of the liquid phase on the global observed reaction rate. This was indicated by the statistically inconclusive predictions made by the three rate expressions. The second item deserving further attention is the indication that nascent
Table V. Product Distribution for DBT Hydrodesulfurization with in Situ Hydrogen Generation run 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21
2,
0.013 72 0.002 31 0.003 99 0.034 48 0.037 78 0.044 55 0.053 45 0.26 12 0.031 15 0.014 65 0.024 72 0.024 08 0.025 80 0.021 95 0.022 21 0.029 87 0.025 97 0.064 24 0.023 54 0.029 46 0.023 53
zco
ZCO*
‘HZ
ZDBT
zBP
ZCHB
ZTHDBT
ZHsS
ZDEC
0.1625 0.1693 0.1719 0.0887 0.0891 0.0925 0.0978 0.0800 0.0724 0.0562 0.0615 0.0654 0.0636 0.0565 0.0589 0.0686 0.0637 0.1096 0.0741 0.0771 0.0687
0.1290 0.1579 0.1938 0.2507 0.2317 0.2073 0.1796 0.2157 0.1879 0.1869 0.1584 0.1967 0.1780 0.1678 0.1537 0.1888 0.1787 0.2127 0.1837 0.1686 0.1577
0.1234 0.1549 0.1870 0.2472 0.2292 0.2049 0.1777 0.2103 0.1830 0.1819 0.1544 0.1855 0.1678 0.1583 0.1455 0.1796 0.1692 0.2096 0.1772 0.1632 0.1531
0.029 15 0.027 08 0.021 75 0.008 33 0.009 21 0.010 57 0.011 75 0.010 60 0.012 38 0.013 38 0.014 77 0.009 79 0.011 20 0.012 38 0.013 60 0.010 62 0.011 22 0.009 78 0.011 81 0.012 79 0.013 97
0.001 722 0.000 913 0.002 000 0.001 112 0.000 808 0.000 801 0.000 655 0.001 887 0.001 669 0.001 610 0.001 377 0.004 027 0.003 679 0.003 401 0.002 936 0.003 325 0.003 463 0.001 114 0.002 627 0.002 224 0.002 075
0.000 191 8 0.000 121 4 0.000 241 0 0.000 148 5 0.000 087 9 0.000 063 2 0.000 070 2 0.000 191 7 0.000 182 6 0.000 195 8 0.000 114 9 0.000 439 0 0.000 382 5 0.000 355 0 0.000 286 3 0.000 370 4 0.000 358 8 0.000 088 8 0.000 158 7 0.000 106 7 0.000 000 0
0.000 607 5 0.000 299 7 0.000 750 6 0.000 232 7 0.000 195 6 0.000 227 2 0.000 160 2 0.000 352 3 0.000 342 2 0.000 398 7 0.000 328 6 0.000 456 4 0.000 451 7 0.000 446 5 0.000 411 7 0.000 363 2 0.000 378 4 0.000 200 2 0.000 248 4 0.000 221 4 0.000 218 5
0.001 913 0.001 035 0.002 241 0.001 267 0.000 895 0.000 864 0.000 725 0.002 079 0.001 852 0.001 806 0.001 492 0.004 466 0.004 061 0.003 756 0.003 222 0.003 696 0.003 822 0.001 203 0.002 785 0.002 331 0.002 075
0.5378 0.4861 0.4163 0.3678 0.4011 0.4382 0.4781 0.4528 0.5091 0.5430 0.5829 0.5092 0.5450 0.5751 0.5992 0.5148 0.5432 0.3916 0.5239 0.5439 0.5786
Ind. Eng. Chem. Process Des. Dev. 1986, 2 5 , 284-289
284
hydrogen seems to be significantly more active than diatomic hydrogen in desulfurizing dibenzothiophene. Because of global depletion of the sweet crude, high sulfur oil is being used more and desulfurization is becoming a major cost factor due to the prohibitive cost of hydrogen. This work indicates that partial combustion flue gases (CO, H,, HzO,some COz, etc.) can be fed directly into the hydrotreater without the necessity of producing hydrogen and separating it. If needed, additional steam can be supplied to the reactor to consume nearly all the CO.
Acknowledgment This work was partially supported by the Center for Energy and Mineral Resources under Grant CEMR-18705. In addition, A. Akgerman is the Shell Faculty Fellow, and Shell's contribution is greatly appreciated.
Nomenclature A = Arrhenius constant C'. = liquid-phase concentration of component J' = vector of component liquid-phase concentrations DEC = decalin E -= activation energy f(C,) = concentration function in rate expression k , = overall rate constant ko = A exp(-E/RTo) ng = molar flow rate of component a n, = initial flow rate of component a number of quadrature points used in the left-hand npl, an right-hand = sides of eq 7, respectively ro = reaction rate for consumption of component a T = absolute temperature To = reference temperature for evaluation of ko and E W = weight of catalyst inside reactor WGS = water gas shift reaction xd = conversion x, = experimentally determined conversion at reactor exit wI = weights of individual quadrature points z = overall mole fraction Registry No. COO,1307-96-6;Moo3, 1313-27-5;HP,1333-74-0; dibenzothiophene, 132-65-0.
4
T
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Received for review August 3, 1984 Revised manuscript received March 4, 1985 Accepted July 3, 1985
Pressure Fluctuations in a Gas-Solid Fluidized Bed with a Screen Y. W. Huang, L. T. Fan,' J. C. Song,+ and N. Yutanlr Department of Chemical Engineering, Durland Hall, Kansas State University, Manhattan, Kansas 66506
To enhance the fundamental understanding of the effect of screen packings on the performance of the gas-solid fluidized bed in terms of pressure fluctuations, the simplest configuration, a single screen installed across the entire section of the bed, was investigated. The effect of the screen insertion on the performance of the bed was determined semiquantitatively by calculating on-line statistical parameters of the pressure fluctuations. The results indicate that the bed stability can be enhanced by insertion of a screen.
The heterogeneity caused by bubbling, slugging, and channeling in a fluidized bed is undesirable because they reduce the extent of fluid-solid contact and, thus, the efficiency of process (Bakker and Heertjes, 1958, 1960). It
* To whom
correspondence should be addressed. Currently in Chenguang Research Institute of Chemical Industry, Fushun County, Sichuan Province, China. On leave from the Department of Chemical Engineering, Tokyo University of Agriculture and Technology,Koganei, Tokyo, Japan 184. t
has been suggested that bubbles, slugs, and channels might be reduced by inserting a screen or other mechanical devices in the bed (Massimilla and Bracale, 1956). Although the screen prevents channeling and slugging, it causes other concomitant effects, e.g., hindering the heat and mass exchanges among various regions in the bed. The effects of screen insertion on the fluidized bed behavior were investigated by different methods. Massimilla and W e s h ~ t e (1960) r carried out a Photographic study ofthe flow pattern in a fluidized bed with screen baffles, and Bailie et al. (1963) used the y-ray attenuation technique
0196-4305/86/1125-0284$01.50/00 1985 American Chemical Society
