Ind. Eng. Chem. Res. 1990,29, 1999-2004
1999
Deactivation of a Ni-Mo/y-A1203Catalyst: Influence of Coke on the Hydroprocessing Activity Fernando Diez and Bruce C. Gates Center for Catalytic Science and Technology, Department of Chemical Engineering, University of Delaware, Newark, Delaware 19716
Jeffrey T. Miller,* Daniel J. Sajkowski, and Simon G. Kukes Amoco Research Center, P.O. Box 3011, Naperville, Illinois 60566
A commercial Ni-Mo/A1203 hydroprocessing catalyst was deactivated with from 1 to 10 w t % coke in operation with either a light catalytic cycle oil (LCCO) or a coal residuum. Coke deposition occurred during the first 20 h. For LCCO, the quantity and aromaticity of the coke as measured by 13C MAS NMR was dependent on the reactor temperature, increasing from 1.5% carbon and 40% aromatic (sp2 carbon) a t 335 "C t o 3.0% carbon and 70% aromatic a t 395 "C. Rate constants for hydrogenation, hydrodesulfurization (HDS), and hydrodenitrogenation (HDN) were approximately 10% lower for the catalyst coked with LCCO a t 335 "C and 20% lower for the catalyst coked with LCCO a t 395 OC than for the coke-free catalyst. A t higher coke levels, however (10.6 w t 7% C with coal residuum), about 75% of the hydrogenation and HDS activity and 95% of the HDN activity were lost. From the N2 pore size distribution, the coke thickness was modeled and estimated to be 0.3 and 0.5 nm for 1.5 and 3.0 w t % C, respectively. T h e coke thickness increased to 1.1nm a t 10.5 wt % carbon. Surface analysis by X-ray photoelectron spectroscopy (XPS) indicated that the alumina support may accumulate coke more rapidly than the molybdenum and nickel sulfides. It is proposed that the loss in the catalyst activity resulted from blockage of the MoS2 crystallite edges (active sites) by coke formed on the alumina support.
Introduction Catalytic hydroprocessing is used to upgrade fossil fuels ranging from light petroleum naphtha to vacuum gas oils and residua. The catalysts are typically sulfided CoOMo03/yA120, (Co-Mo) or Ni0-Mo03/y-A1203 (Ni-Mo), which can undergo deactivation a t widely different rates depending on the feed and processing conditions. A naphtha hydrodesulfurization catalyst may remain active for years, whereas a residuum hydroprocessing catalyst may last only a few months. Catalyst deactivation may result from (1) chemical poisoning, as feed components adsorb on catalytic sites or react to form other strongly adsorbed species, (2) physical blockage of the catalyst surface, preventing access of reactants to the catalytic sites, or (3) physical blockage of catalyst pores (or pore mouths), preventing access of reactants to the interior pore volume. In commercial hydroprocessing applications, the temperature of the catalyst in a fixed-bed reactor is continually raised to compensate for the deactivation, with some product properties (such as the sulfur content) held constant and providing a criterion for the temperature increase. When the average reactor temperature required to maintain performance is plotted against time on stream, the typical result is a S-shaped curve. For example, in petroleum residuum hydroprocessing (Quann et al., 1988), there is, typically, an initial rapid temperature increase over a period of about 10 days, followed by a nearly linear temperature increase over the next several months and then a rapid, accelerating temperature increased at the end of the run. The rapid initial deactivation has been attributed to chemical poisoning by coke, but some authors have attributed it instead to chemical poisoning by nickel and vanadium sulfides (Tamm et al., 1981). The slow deactivation period following the initial rapid deactivation has been associated with buildup of layers of coke and metal sulfides, and the final catastrophic deactivation has *Author to whom correspondence should be addressed.-
been associated with pore mouth blockage by metal sulfides (Tamm et al., 1981; Dautzenberg and DeDeken, 1984; Quann et al., 1988; Pereira et al., 1987). Although this pattern of catalyst deactivation is well established for many applications, the competing deactivation processes have not been well resolved. The greatest attention has been paid to catalysts used in heavy residuum upgrading, since catalyst cost is a significant fraction of processing costs in these applications, whereas it is not in distillate hydroprocessing. Deactivation resulting from demetallation is relatively well understood (Quann et al., 1988; Stohl et al. 1987; Stohl and Stephens, 1987), and consequently, one of the most important unanswered questions concerns the deactivation by coke (Ternan et al., 1979; Ramachandran and Massoth, 1982). The objectives of this research were to prepare hydroprocessing catalysts with various coke contents (in the absence of metal sulfide deposits) and to determine the nature of the coke deposits and their effect on the catalyst performance. The catalysts and coke were characterized by physical methods, including elemental analysis, I3C CP MAS NMR spectroscopy, BET surface area measurements, pore size distribution determined by N2desorption, and X-ray photoelectron spectroscopy (XPS). Fresh and coked catalysts were evaluated in a flow microreactor with a synthetic feed mixture containing phenanthrene, dibenzothiophene, and quinoline in order to determine how the coke affects the activity for aromatic ring hydrogenation, hydrodesulfurization (HDS), a n d hydrodenitrogenation (HDN), respectively. The reactant mixture gave simple enough product mixtures to allow precise product analysis and estimation of rate constants for the individual reactions in the hydrogenation, hydrodesulfurization, and hydrodenitrogenation networks. The rate constants could then be related to the composition and properties of the catalysts.
Experimental Methods Catalyst and Feeds. The catalyst was American Cyanamid HDS-SA, a commercial alumina-supported
0888-588519012629-1999$02.50/0 0 1990 American Chemical Society
2000 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 Table I. Properties of HDS-SA Catalyst Samples sample treatment fresh (unsulfided) used LCCO/335 OC/20 h used LCCO/335 OC/iO h used LCCO/395 "C/20 h used LCCO/395 'C/20 h followed by 335 'C/52 h used coal residuum/72 h Per gram of catalyst plus carbon.
m2/g 146 131 125 129 129 113
SRET,(I
Per gram of fresh catalyst.
Table 11. Feed Properties
specific gravity, "API elemental anal.
c, wt 70 H, wt % s, wt %
0 , W t 70 N, PPm hydrocarbon type, wt % paraffins total aromatics naphthalenes phenanthrenes ramscarbon asphaltenes (toluene soluble. hexane insoluble) preasphaltenes (tetrahydrofuran soluble, toluene insoluble)
feed panasol-coal LCCO residuum blend 21.9 -1.1 89.58 10.37 0.55 540
90.56 7.48 0.21 1.18 5840
30.5 69.5 25.4 5.2 21.6 15.3 15.3
nickel-molybdenum hydroprocessing catalyst (NiO, 3.1 wt %; Moo3, 18.3 wt W ) supplied as 1/!6-in. extrudates. Catalyst physical properties are listed in Table I. Light catalytic cycle oil (LCCO) was used to deposit small amounts of coke on the catalyst. LCCO is a middle distillate refinery product generated in fluid catalytic cracking units; properties of the LCCO are given in Table 11. The LCCO is highly aromatic (containing mostly two-ring compounds), with a boiling range of 190-380 "C. It was chosen because it contained low concentrations of heteroatoms such as sulfur (0.5 w t % ) and nitrogen (500 ppm) and negligible metal contaminants. The concentration of heavy polyaromatics in LCCO is low; the content of four-ring aromatics was 0.1 wt %, and there was a negligible concentration of five-ring aromatics. A second, heavier feedstock was used to prepare a coked catalyst with a higher carbon content. The feed was a 50:50 blend by weight of deashed coal residuum (Wilsonville, Al, run 247, Illinois No. 6 coal) with panasol, a byproduct of naphtha reforming; panasol consists primarily of alkyl-substituted naphthalenes. Metal contaminants in the deashed coal residuum were low, less than 1 ppm for Ni, V, and Ti and approximately 200 ppm Fe. Properties of the panasol-coal residuum mixture are given in Table 11. Coke Deposition with LCCO and Coal Residuum. Coking of the pre-sulfided catalyst with LCCO was conducted in a standard bench-scale high-pressure, trickel-bed flow reactor. The WHSV (g of gas oil feed/(h.g of catalyst)) was 5.0, the hydrogen flow rate was 10 L / h (STP), and the reactor pressure was 85 atm. Two different temperatures were used (335 and 395 "C) to determine the effect on coke formation. Coke deposition was also conducted for two different periods of reactor operation (20 and 72 h) at each temperature. An additional catalyst was prepared by first depositing coke from LCCO a t 395 "C for 20 h, followed by reaction at 335 "C for an additional 52 h. The coked catalysts were Soxhlet extracted for 2 days with toluene followed by 1 day with cyclohexane and
m2/g PV? cm3/g PV,b cm3/g 146 0.45 0.45 133 0.42 0.43 127 0.40 0.41 133 0.36 0.37 133 0.38 0.39 126 0.33 0.37
&ET,*
C, w t % 0.07 1.7 1.6 3.0 2.7 10.6
S,wt
%
6.09 6.10 5.90 6.05 ND'
ND = not determined.
then dried at 110 "C under vacuum overnight. An additional catalyst was prepared by depositing coke from coal residuum at 415 "C and 0.67 WHSV for 72 h. The coal residuum was pumped to a continuous-flow, stirred autoclave reactor containing 15 g of catalyst and operated at 135 atm and a H2 flow rate of 13 L/h (STP). The reactor was cooled to room temperature and the catalyst removed and Soxhlet extracted overnight with tetrahydrofuran and then dried at 100 "C under vacuum. Catalyst Analysis. Analyses for elemental carbon and sulfur were performed by the Amoco Oil Analytical department. XPS measurements were obtained with a Hewlett-Packard 5905B X-ray photoelectron spectrometer using monochromatic A1 K a radiation (hu = 1486.6 eV). Surface compositions were calculated in the usual manner by correcting for instrumental parameters, photoionization cross sections, and escape depths. I3C CP MAS NMR spectra were obtained on a Joel FX6OQS Fourier transform spectrometer equipped with a chemognetic solid accessory. Spectra were obtained by using a 8-ps lH 90" pulse, a 2-ms contact time, and a relaxation delay of 0.5 s. The spin rate of 2.4 kHz, and chemical shifts were referenced to tetramethylsilane. The fresh and coke catalysts were characterized by N2 physisorption/desorption using a Micrometrics Corporation Digisorb Model 2500 to obtain BET surface areas (BET,s), pore volumes (PVs), and pore size distributions (PSDs). Microreactor Experiments. Activities of the catalysts that had been previously coked with either LCCO or the coal residuum were evaluated in flow microreactor experiments with a mixture containing 2.5 mmol/L phenanthrene (Aldrich, purity 98+%), 0.37 mmol/L dibenzothiophene (Eastman, 98%), 1.87 mmol/L quinoline (Aldrich, 99%), and 0.45 mmol/L n-dodecane (Aldrich, 99+%), dissolved in cyclohexane (Aldrich, 99.9%); these concentrations were measured at room temperature. n-Dodecane was used as an internal standard. The feed also contained 0.1 wt % carbon disulfide to maintain the catalyst in a sulfided state. The catalyst was diluted with inert 90-mesh a-alumina. Each catalyst charge was pre-sulfided with a flowing mixture of 10% H2S and 90% H2 (Matheson) for 2 h at 400 "C and atmospheric pressure. The high-pressure reactor was a 316 stainless steel cylinder, 25.4 cm in length and 0.32 cm in internal diameter, mounted in a heating block. The liquid feed to the reactor was presaturated with hydrogen at 135 atm in two high-pressure 1-L autoclaves (Autoclave Engineers) equipped with mechanical stirrers. It has been estimated that the equilibrium mole fraction of hydrogen in cyclohexane is approximately 0.05, and therefore, the hydrogen was in stoichiometric excess relative to the organic reactant (Girgis, 1988). Under typical reaction conditions, the hydrogen concentration at the reactor outlet was a t most a few percent less than that at the inlet. The liquid feed was pumped downward through the reactor at 163 atm by a high-pressure liquid chromatograph pump (Waters M-6000 A), and the product was collected in a 500-cm3 high-pressure receiver. The reaction equipment is de-
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2001
".".
0.07
I
0.06 0.05
3 I g
0.04
0.03 0.02
0.01
0.00
Pore radius, nm Figure 1. Pore size distribution of the fresh and used catalysts.
scribed in more detail elsewhere (Girgis, 1988). The advantage of the design of the microreactor is that only two phases were present in the microreactor, the solid catalyst and the liquid reactant, because the hydrogen is dissolved in the feed. Ideal plug flow was nearly realized with this system. The reactor was charged with 0.04 g of 80-100-mesh particles mixed with 0.3 g of a-alumina. The mixture was placed in the central part of the reactor, with the upper and lower sections filled with alumina particles. Under these conditions, there was no significant mass-transfer limitations (Girgis, 1988).
Results and Discussion Characterization of Catalysts. Elemental analysis showing carbon on coked catalysts are presented in Table I. The catalysts used for hydrotreating LCCO a t 335 "C contained 1.6 wt 90C, and those used at 395 "C contained 3.0 wt % C. There was a neglegile difference in the amount of coke formed between 20 and 72 h. In contrast, the catalyst used at 415 "C with coal residuum for 72 h contained 10.6 wt % C. BET surface areas of the catalysts are summarized in Table I. Pore size distributions are presented in Figure 1. There was a slight decrease in surface area, from 146 m2/g for the fresh catalyst to 130 m2/g for the coked catalysts, in operation with LCCO. The pore volume also decreased slightly from 0.45 to about 0.40 cm3/g. The different temperatures of coking with LCCO evidently did not lead to significant differences in surface area and pore volume. The surface area and pore volume of the catalyst coked in operation with the coal residuum showed the greatest decrease, to 113 m2/g and 0.33 cm3/g, respectively. Estimates of the thickness of the coke have been made with the assumption that the coke layers are uniform and the catalyst pores are cylindrical and uniform. For example, the pore size distribution for the catalyst coked with LCCO at 395 "C and containing 3.0 w t % C was calculated from the pore size distribution of the fresh catalyst, assuming that the coke thickness was 0.5 nm. The agreement between the observed and calculated pore size distributions is good, Figure 2, considering the simplicity of the model. Similarly, the pore size distribution has been calculated for the catalyst coked with the coal residuum and containing 10.6 w t % C, assuming that the coke thickness was 1.1nm. Again, the agreement with the observation is good, Figure 3. The results are consistent with the conclusion that the deposits were nearly uniform and increasing in thickness with increasing amounts of coke. The aromaticity of the coke, as measured by the percentage of sp2 carbon in the 13CCP MAS NMR, Table I11 and Figure 4, depends strongly on the reaction temperature, increasing from 40% at 335 "C to 70% at 395 "C with
0
2
4
6
0
10
Radius. nm
Figure 2. Pore size distribution measured for the catalyst used in hydroprocessing of LCCO at 395 OC, containing 3.0 w t % C, and pore size distribution calculated for that sample assuming that the coke deposit was uniform, with a thickness of 0.5 nM.
::I
0'07
1
,$? 0.04
_0
2
4
6
0
lo
Radius, nm
Figure 3. Pore size distribution measured for the catalyst used in hydroprocessing of coal residuum, containing 10.6 wt % C, and pore size distribution calculated for that sample assuming that the coke deposit was uniform, with a thickness of 1.1 nM.
LCC0/3359:
0
s?
0
a
2002 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 Table 111. Comparison of Catalysts by 13C NMR SDectroscoDv wt % sample run time % C fresh catalystn 0.5 used catalyst LCCO/335 "C 20 1.7 LCCO/335 "C 70 1.6 LCCO/395 "C 20 3.0 LCCO/395 "C/335 "C 20/52 2.7 coal residuum/415 "C 72 10.6
A,) Phenanthrene hydrogenatbn
r
7 - 7
Dlhydrophenanthrene
aromatic 13c
32b 43 39 73 70 73
Sulfided catalyst was Soxhlet extracted similar to catalyst coked with LCCO prior to NMR analysis. bPrecise determination of percent aromatic carbon was limited due to the low carbon content of the catalyst.
LCCO. The amount (2.7 wt % C) and aromaticity (70 sp2) of the coke formed on the catalyst treated first with LCCO at 395 and then a t 335 "C are characteristic of the operation at the higher temperature. The aromaticity of the coke formed from the coal residuum at 415 "C is about the same as that of the coke formed from the much lower boiling LCCO feed a t 395 "C; there is evidently not a correlation between the concentration of heavy polynuclear aromatics in the feedstock and the aromaticity of the coke formed on the catalyst surface. The XPS data, Table IV, show that the surface Mo concentration of the fresh sulfided (extracted) catalyst was approximately 2.0 atom 70.With the first small amount of coke deposited on the catalyst surface, this concentration decreased slightly, to 1.7 atom %. With further accumulation of coke, the Mo concentration decreased by less than 10% even for the catalyst coked by the coal residuum, which contained 10.6 wt % C. The Ni/Mo atomic ratio was constant at 0.3 in all the samples. In contrast, the surface A1 concentration decreased sharply as the coke content increased. (The concentration of surface carbon on the fresh, unsulfided catalyst is associated with residual carbon in the vacuum chamber.) The A1 concentration was about 26.4 atom 70for the catalyst with 1.7 wt % coke and about 19.1 atom % for that with 10.6 w t 7'0 coke. These results are consistent with an earlier report (Fleisch et al., 1984) of increases in the surface Mo/Al ratio with increasing coke on the catalyst. XPS data suggest that the coke was deposited preferentially on the alumina support rather than on the catalytically active sulfide components. Possibly the activity of the latter for hydrogenation helps to keep the surface preferentially clean by removal of coke or coke precursors. Characterization of Catalysts by Model Compound Reactions. Experiments with fresh and coked catalysts (LCCO/335 "C, LCCO/395 "C/335 "C, and coal residuum/415 "C containing 1.7, 2.7, and 10.6 wt % C, respectively) were carried out with the flow microreactor. The reaction temperature was 350 "C and the pressure 163 atm. The reaction networks characterizing hydroprocessing of phenanthrene, quinoline, and dibenzothiophene in mixtures similar to those used in this work have been investigated in detail (Girgis, 1988). Simplified reaction networks presented in Figure 5 have been assumed to represent the data obtained in this work. Each reaction is
B.) Dibenzothiophene hydrodeaulfurlzation
Dibenzothiophene
Biphenyl
C.) Quinoline hydrodenitrogenation
Quinoline
5.6.?&tetrahydroquinoline
123.4-tetrahydroQJlnOHne
0-Propylaniline
Decahydro-
wholine
Figure 5. Simplified reaction networks for hydroprocessing of phenanthrene, dibenzothiophene, and quinoline (Girgis, 1988).
approximated as pseudo first order in the reactant. Phenanthrene Hydrogenation. At the low conversions, only two products, 9,lO-dihydrophenanthreneand 1,2,3,4-tetrahydrophenanthrene, were found, Figure 5A. When the data were fitted to a network in which these reactions were considered reversible, negative rate constants were calculated for the dehydrogenation reactions; consequently, the hydrogenation was considered to be irreversible. Dibenzothiophene Hydrodesulfurization. Biphenyl was the only product, and the simple network of Figure 5B was, therefore, assumed. Quinoline Hydrodenitrogenation. The simplified reaction network for hydroprocessing of quinoline is presented in Figure 5C. Quinoline rapidly undergoes partial hydrogenation to give tetrahydroquinolines. Carbon-nitrogen bond cleavage in 1,2,3,4-tetrahydroquinoline leads to o-propylaniline, which reacts to give hydrocarbons (propylbenzene and propylcyclohexane) plus ammonia. Only small amounts of decahydroquinoline were observed. The concentration of each of these reaction intermediates and products (except ammonia and HzO) was measured at each feed rate. Pseudo-first-order rate constants for the reactions in each of the simplified networks were calculated by minimizing the least-squares error for the data set. Each datum was weighted equally in the computation of the rate constants for each network. The first-order rate constants for each catalyst for the phenanthrene, dibenzothiophene, and quinoline reaction networks are summarized in Table V. Figures 6-8 show typical fits of the experimental results; the curves were
Table IV. Surface Compositions of Catalyst (Atom % ) Estimated by XPS comDosition sample P Mo A1 Ni 0 c fresh (oxide) 1.6 2.4 31.2 0.75 56.2 7.8 1.2 2.0 28.7 0.7 52.7 fresh (sulfided, extracted) 11.4 LCC0/335 "C 1.2 1.7 26.4 0.6 52.9 13.0 LCCO/395 OC 1.8 1.8 23.6 0.5 47.4 20.5 coal residuum 1.1 1.6 19.1 0.5 40.5 34.5
S 3.3 4.0 4.4 2.7
Mo/Al 0.077 0.069 0.064 0.076 0.084
Ni/Mo 0.31 0.35 0.36 0.28 0.31
Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990 2003 Table V. Activities of Fresh and Deactivated Catalysts at 350 "C and 163 atm catalyst,' wt ?' % C LCCO/ LCCO/395 coal 335 'C OC/335 OC residuum fresh (1.7 wt W ) (2.7 wt %) (10.6 wt 70) First-Order Rate Constant in Hydroprocessing Networkb phenanthrene kl 0.073 0.066 0.055 0.019 k2 0.050 0.042 0.040 0.014 dibenzothiophene k 0.120 0.13 0.086 0.004 quinoline kl 8.0 9.1 8.2 10.9 k2 2.4 2.2 2.4 3.3 k3 3.8 3.3 3.2 0.80 k4 0.99 0.78 0.95 0.26 k5 2.7 2.6 2.1 0.06 k6 2.0 2.0 2.2 C a Coke deposited under conditions given in Table I and in text. *In units of L/(g of catalyst-h), where the mass of catalyst refers to the fresh unsulfided form. Value could not be determined precisely.
0
2
1
g of catalyst x h
mmol of Quinoline
+
Hydrocarbon products + decahydroq*noline 12,3,4-Tetrahydroquiiine Whohe o-Propylan%ne A 5.6.78-Tetrahydroquinohe
A
Figure 8. Hydroprocessing of quinoline in the feed containing phenanthrene and dibenzothiophene: The reaction products. The catalyst had been coked in operation with LCCO at 335 " C and contained 1.7 wt % C. The curves are the predictions of the reaction network shown in Figure 5 with the parameter values given in Table V.
0.2
01
0 0
0.0
1.o
0.5
Q of catalyst x h mmol of Phenanthrene
Dihydrophenanthrene, fresh catalyst 0 Tetrahydrophenanthrene. fresh catalyst A
A
Dihydrophenanthrene, catalyst coked with coal residuum Tetrahydrophenanthrene, catalyst coked with coal residuum
Figure 6. Hydrogenation of phenanthrene in the feed containing dibenzothiophene and quinoline. The data show a comparison of the fresh catalyst and the coked in operation with a coal residuum (Table I). The curves are the predictions of the reaction network shown in Figure 5 with the parameter values of Table V. 0.5
I 0
I
2
I
6
S A
I
8
1
0
Wt% carbon on catalyst x k Dbenzothiophene hydrodesulfurization
f, Li }
0
I
4
1
2
3
4
5
6
7
8
9
o of catalyst x h mmol of dibenzothiophene Fresh catalyst o Catalyst coked with coal residuum (10.6% carbon)
Figure 7. Hydrodesulfurization of dibenzothiophene in a mixture with phenanthrene and quinoline. The data show a comparison of the fresh catalyst and a catalyst coked in operation with a coal residuum (Table I). The lines are the predictions of the reaction network shown in Figure 5 with the parameter values given in Table V.
determined from the calculated rate constants. Results are shown for the fresh catalyst and that coked by coal residuum for phenanthrene hydrogenation, Figure 6, and dibenzothiophene hydrodesulfurization, Figure 7. The data are well represented by the simplified reaction net-
Phenanthrene hydrogenation
Figure 9. Effect of coke deposition of the activity of a sulfided Ni-Mo/y-A120, catalyst for hydrogenation, hydrodesulfurization, and hydrodenitrogenation. The rate constants are for the reaction networks shown in Figure 5.
works. For quinoline, six rate constants were calculated which best fit the data set. For example, the results for quinoline conversion with the catalyst containing 1.7 wt % coke (LCCO/335 "C)are shown in Figure 8. Because of the more complex reaction network for quinoline, estimated rate constants are less precise than for the reactions of phenanthene and dibenzothiophere. For the most part, the product distribution was unaffected by the coke. The rates of hydrogenation and hydrodesulfurization decreased roughly in proportion with increasing coke. The relative rate constants, normalized to the values for the fresh catalyst, are shown as a function of catalyst carbon content in Figure 9. The data are
2004 Ind. Eng. Chem. Res., Vol. 29, No. 10, 1990
scattered, but they show a pattern of decreasing rate constants with increasing carbon content of the catalyst. Deposition of 3 wt % C resulted in a loss of about 20% of the fresh activity for each of the hydroprocessing reactions. At the highest coke concentration, deposition of 10.6 wt % C resulted in a loss of 75% of fresh activity for hydrogenation and hydrodesulfurization. Hydrodenitrogenation activity, however, was markedly reduced. The activity of the catalyst containing 10.6 wt % C was less than 5% of the value measured for fresh catalyst as indicated by the value of k5 for the quinoline network. Consequently, the value of k6 could not be determined with any certainty, Table V. Further, since the hydrogenation-dehydrogenation reactions characterized by k, and k, in the quinoline network were fast (the reaction was near equilibrium), the values of these parameters could not be determined precisely. Although heavily coked catalysts still have activity for quinoline hydrogenation, albeit reduced. they no longer have much activity to cleave the C-N bond to form NH,. Processing Implications. The results of this research have implications for catalytic hydroprocessing. They indicate that the carbon deposition on nickel sulfidemolybdenum sulfide hydroprocessing catalysts occurs rapidly during the first 20 h of operation. Subsequent coke deposition at the same reaction temperature evidently occurs much more slowly, likely over a period of months. With an increase in reaction temperature, more coke is deposited rapidly. Returning to a lower temperature does not immediately reduce the amount of coke or its aromaticity. Since higher concentrations of coke led to decreased catalytic activity, it is inferred that even a brief increase in reaction temperature will result in almost immediate activity loss. A similar result might be expected from a change in the feed composition (at constant operating conditions), for example, a change from LCCO to coal residuum. Returning to the original feed would not be expected to immediately restore the catalyst activity.
Conclusion Initial coke formation on surface of hydroprocessing catalysts used with light catalytic cycle oil and coal residuum is rapid, occurring during the first 20 h of operation. For a given feed, the amount of aromaticity of the coke are determined, primarily, by the maximum reaction temperature. The coke layer is estimated to be approximately 0.5 nm in thickness when the catalyst contains 3 wt % C (LCCO/395 "C) and 1.1nm in thickness when the catalyst contains 11wt % C (coal residuum/415 "C). The accuracy of the calculated pore size distributions and the simplicity of the coke and pore size distribution models, however, make these values only approximate. Although the whole catalyst surface accumulates coke, the alumina appears to accumulate it more rapidly than the catalytically active nickel and molybdenum sulfides. In the catalyst containing 10.6 wt % C, the exposure of the nickel and molybdenum sulfides measured by XPS was not much different from that in the catalyst containing 3.0 wt % C. The catalyst containing 10.6 wt % C, however, had lost most of its hydroprocessing activity. We speculate that the active nickel and molybdenum phase is maintained relatively coke free as a result of self-cleaning process but that activity is lost as coke accumulates on the alumina support, growing in thickness. The loss of catalytic activity may result from edge blockage of the Ni-promoted MoSz crystallites. The active sites of the catalyst are believed to be present a t the edge of these crystallites (Roxlo et al., 1986; Topsoe et al., 1986). Since the MoSz crystallites are estimated to be 0.615-nm thick with lateral dimensions of
about 1.0-1.5 nm (Clausen et al., 1981; Candia et al., 1984; Bouwens et al., 1988; Hayden and Dumesic, 1987; Hayden et al., 1987), coke deposits with a thickness of less than about 0.5 nm might be relatively ineffective in blocking catalytic sites. As the coke thickness increases beyond about 0.7 or 0.8 nm, however, catalytic activities might be expected to decrease rapidly (Hayden and Dumesic, 1987). A conceptually similar model for catalyst deactivation has recently been proposed (Yamamoto et al., 1988).
Acknowledgment We thank the Foundation for the Support of Research in Applied ScienGe and Technology in Asturias, Spain, for support of Dr. Diez and Amoco Oil Company for support of the research.
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Received for review December 18, 1989 Revised manuscript received J u n e 5, 1990 Accepted June 25, 1990
