Coping with Catalyst Deactivation in Hydrocracking: Catalyst and

3354

Ind. Eng. Chem. Res. 1997, 36, 3354-3359

Coping with Catalyst Deactivation in Hydrocracking: Catalyst and Process Development J. W. Gosselink* and W. H. J. Stork Shell International Oil Products B. V., Shell Research and Technology Centre, Amsterdam, Badhuisweg 3, 1031 CM Amsterdam, The Netherlands

During operation, hydroprocessing catalysts undergo deactivation by coke formation, metals deposition, poisoning, etc. In the literature catalyst deactivation has been extensively studied and reviewed, with focus on the description and understanding of deactivation phenomena. For oil refinery process and catalyst development, the maximum catalyst life, or cycle length, is of prime importance. The catalyst cycle length is determined by initial activities for the relevant reactions, their decline rates, changes in both product yields and qualities during deactivation, and (equipment) fouling by byproducts. The objectives in the development of more stable catalysts are longer cycle lengths in the refineries, intake of heavier feedstocks, and increase of severity of operation. Two processes, hydrocracking and mild hydrocracking, will be highlighted to illustrate how catalyst deactivation is coped with in an industrial environment. Introduction

Coke Deactivation and Poisoning of Hydroprocessing Catalysts

During operation, hydroprocessing catalysts undergo deactivation by coke formation, metals deposition, poisoning, etc. In the literature catalyst deactivation has been extensively studied and reviewed [1]. Much attention has been given to the area of hydroconversion (particularly residue hydroconversion) [2]. In general, however, these reviews have focused on the description and understanding of deactivation, with little consideration to process or catalyst development. Recently, performance testing of hydroconversion catalysts was critically reviewed [3]. For oil refinery process and catalyst development, the maximum catalyst life, or cycle length, is of prime importance [3]. Increasing cycle length reduces the number of catalyst replacements and consequent downtime of the unit. However, increasing severity of operation while maintaining cycle length is also a valuable option. Advantages are, for instance, more conversion in the unit (mild hydrocracking) or processing of heavier and less costly feedstock (hydrocracking). This is obvious when one considers that most catalysts are regenerable and the availability of presulfurized catalysts and offsite regeneration often allows shorter downtimes of the processing units. The cycle length is determined by [3] (1) the initial activities of the catalysts for the relevant reactions and their deactivation rates in relation to the design endof-run temperature limits, (2) changes of product selectivities with time on stream in relation to bottlenecks in the fractionator section and economic considerations, (3) changes in product qualities with time on stream in relation to product specifications and economic considerations, and (4) fouling of equipment [4]. The present paper focuses on the deactivation of hydroprocessing, in particular (mild) hydrocracking, catalysts and coping with it in terms of catalyst development, process development, and modeling. For the purpose of this paper a relatively simple kinetic model is used to illustrate catalyst deactivation phenomena. A more complete description of all hydroprocessing phenomena is reported elsewhere [5]. S0888-5885(96)00599-4 CCC: $14.00

It is very difficult to describe and model deactivation in oil refinery processes as a result of the complex multicomponent systems in the oil streams. For example, simple gas oil hydrodesulfurization is described by continuous distribution functions for both reaction rate constants and effective pore diffusion coefficients [6], illustrating the problems to be encountered with complex hydroconversion processing where multifunctional catalysts are applied. When restricted to (mild) hydrocracking virtually metals-free feedstocks, irreversible deactivation is the result of coke deposition on the catalysts. Coke is formed from precursors in the feed (asphalthene-like molecules) [7, 8] or from precursors generated through condensation reactions as part of the process itself (polynuclear aromatics) [4, 9]. Coke deposition on hydroprocessing catalysts should be considered at several levels: on an active surface, the metal sulfide crystallites (or in general the hydrogenating function) clear an annular zone of alumina surface around the metal crystallites from carbon. [10] Outside the annulus there is a uniformly thick layer of coke whereas inside the annulus no coke is present. (This model [10] is schematically illustrated in Figure 1.) In contrast, at extreme hydroprocessing conditions though, viz., high temperature and low hydrogen pressure, dehydrogenation is favored over hydrogenation [9]. At the level of the pores, with increasing carbon content pore mouths become plugged, leading to an increase in the number of active sites apparently poisoned (by coke), which is larger than the actual number of poisoned sites [11]. Eventually, at a high level of coke deposition, the catalyst particle or crystallite exhibits areas of unpoisoned (by coke) sites, which are nevertheless not accessible. This phenomenon is well described by stochastic modeling and percolation theories [12, 13] and is obviously dependent on the pore structure [14, 15]: e.g., pore sizes, cavities, and type of pore network. In addition to coking, the catalysts in (mild) hydrocracking suffer, the one more than the other and depending on process conditions, from poisoning of the © 1997 American Chemical Society

Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 3355

Figure 1. Metal sulfide crystallites clearing annulus on the surface.[10]

active sites by other feed molecules, already leading to activity loss at low levels of coke deposition [11]; e.g., zeolite catalysts suffer from poisoning by organic nitrogen containing molecules. This latter poisoning is due to an equilibrium adsorption to active sites and is in principle completely reversible. Nevertheless, it may take a long time to achieve total equilibrium adsorption over the entire catalyst bed, because the concentration of the poisons in the feedstocks is low or the adsorption/ desorption processes are slow, especially at low temperatures, e.g., the second stage of a hydrocracker unit running under low ammonia and low organic nitrogen levels, where it may take some 2000 h to reach a steady state [16]. Consequently, during the initial phase of the catalyst cycle, such poisoning effects may obscure the coke deactivation, making the distinction between both processes difficult. Interestingly, a catalyst not yet poisoned to its equilibrium state may be very susceptible to coking. In commercial practice, catalyst deactivation is generally coped with by increasing the reaction temperature, expressed as °C (F) per 1000 runhours. This experimental measure is related to the fundamental deactivation parameters through the apparent activation energy. A high activation energy results in a lower temperature increase requirement for a given fractional deactivation. Strongly poisoned systems often have high apparent activation energies, since the degree of poisoning decreases with increasing temperature. Thus in a practical application a strongly poisoned system can have excellent long-term stability (although not always good short-term reactor temperature stability). Two processes, hydrocracking and mild hydrocracking, will be highlighted to illustrate how catalyst deactivation is coped with in an industrial environment. The area of (mild) hydrocracking is particularly complex, because it involves both heteroatom removal and cracking, multiple bifunctional catalyst systems are generally used, and product recycle is often applied. Both base functions of the bifunctional catalysts, hydrogenation and cracking, may be affected differently as the catalysts are being deactivated by coke deposition, in particular the cracking function being more affected [17]. This will influence the product yields and qualities with time on stream. Unique to hydrocracking is the application of a liquid recycle feed in combination with full conversion of the feedstock. In the particular case of zeolitic catalysts, size exclusion effects can lead to the build up in the recycle loop of bulky and inert compounds, requiring an increase of operation temperature to maintain overall conversion and resulting in an overall deterioration of the selectivity due to over-

Figure 2. Improvement of the resistance of MHC catalysts to coke deactivation.[19] (A) Conversion and HDS decline rate as function of temperature for conventional HT catalysts. (B) HDS decline rate: improved MHC catalyst versus conventional HT catalyst.

cracking of feed molecules. Furthermore, large polynuclear aromatics lead to both catalyst deactivation and fouling of downstream equipment (heat exchangers, socalled “red-death”), due to their low solubility [4, 18]. Mild Hydrocracking Coping with catalyst deactivation in oil refinery processes is nicely illustrated by the development of mild hydrocracking (MHC) processes and catalysts. MHC imparts extra cracking capacity to a vacuum gas desulfurization (HDS) unit simply by increasing the temperature (increase of severity) [19-24]. Originally, conventional hydrotreating (HT) catalysts were used [25,26]. However, the resulting exponential increase of catalyst deactivation by coke deposition limits the conversion levels that are achievable (see Figure 2). The high temperature and relatively low hydrogen partial pressure in the MHC process favor hydrogen removal reactions and thus coke formation [19]. Obviously, space velocity and feedstock quality also govern the catalyst deactivation [27]. Analyses of spent catalyst from MHC pilot plant runs show an increasing carbon level on catalyst on progression down the reactor, which is caused by a significant decrease of hydrogen partial pressure, again on progression down the reactor. Decrease of the hydrogen partial pressure is induced by both the hydrogenolysis and cracking reactions and the formation of more volatile products. The first step to overcome the limits on the achievable cracking conversion levels was to develop a catalyst that is more resistant to coke deactivation by optimizing both porous texture and metals loading [19]. Figure 2 shows the performance of such an improved MHC catalyst. Because of its higher resistance to coke deactivation, this catalyst can be used at higher operating temperatures, and thus higher cracking levels, and still reach required cycle length and sulfur specifications. A direct increase in cracking activity can be achieved by using multiple catalyst bed systems which comprise a hydrotreating catalyst on top of a zeolite-based hydrocracking (HC) catalyst [20, 28-30]. The hydrotreating catalyst furnishes the necessary HDS and hydrodenitrogenation (HDN) activity. The former is required to meet the sulfur specifications in the feed, while the latter is required to reduce the level of nitrogen contain-

3356 Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997

Figure 5. Improvement of HDS function of zeolitic MHC catalysts.

Figure 3. Poisoning of Z-703 by N-containing molecules at MHC conditions.[19]

catalysts: lower HDS activity and middle distillate (MD) selectivity. Increasing the MD selectivity of the catalyst-stacked beds could be accomplished by replacing Z-753 by more MD-selective zeolite-containing MHC catalysts, while retaining the high resistance to deactivation [20]. Increasing the necessary high HDS activity was achieved by introducing additional dedicated HDS functions during the preparation of the zeolitecontaining MHC catalysts, as was confirmed by standard gas oil HDS tests (see Figure 5). In short, to boost the cracking activity in a MHC reactor, a hydrocracking catalyst should be used, preferably together with a high-stability hydrotreating catalyst. Hydrocracking

Figure 4. Boosting of cracking activity by introducing Z-753. (A) >370 °C conversion and HDS level as function of temperature for C-424 and C-424/Z-753. (B) >370 °C conversion versus HDS level at equal HDS decline rate: C-344/Z-753 versus improved MHC catalyst.

ing organic molecules, which significantly poison the HC catalyst, and thereby enable this catalyst to boost the cracking activity. Figure 3 illustrates the effect of the level of nitrogen containing organic molecules in the total liquid effluent from the hydrotreating (C-424) catalyst on the cracking activity (>370 °C conversion level) of the hydrocracking (Z-703) catalyst at fixed temperature and pressure conditions. By computer modeling of this separation of the activity over the different catalysts, the catalyst volume ratio and bed temperatures could be fine-tuned [20]. Pilot plant testing of these optimised multiple-bed systems with the application of a second stage zeolite containing HC catalyst Z-753 confirmed the increase of cracking activity at equal deactivation rate. Figure 4 shows that at higher temperatures the C-424/Z-753 catalyst stacked bed provides more conversion capacity than a conventional single C-424 catalyst (total catalyst volume of the C-424/Z-753 stacked bed equals that of the C-424 single bed), without significantly affecting the HDS activity. Furthermore, when operating at an almost equal decline rate of the HDS activity under the more difficult low pressure MHC conditions, the C-344/ Z-753 catalyst-stacked bed also provides more conversion capacity than the improved MHC catalyst (see also Figure 2), though at the cost of HDS activity. In general, two serious drawbacks were encountered with the introduction of the zeolitic hydrocracking

Hydrocracking is a very flexible process for the conversion of, among others, flashed distillates, cycle oils and deasphalted oils, and syn-fuels into premium quality gas oil and kerosene and naphtha for reformers. The process can be carried out for full conversion to transportation fuels or for the by-production of hydrowax, an excellent feedstock for, e.g., an ethylene plant or a catalytic cracker. Several modes of operations are being used [31]. In single-stage operation the feedstock is converted in one stage either to produce mainly transportation fuels or partially to produce both transportation fuels and hydrowax. In both the two-stage and series flow modes, the feedstock is treated in the first stage to remove heteroatoms and the hydrogenate aromatics, while this is accompanied by partially cracking to lighter products [32]. The HDN activity of the first stage catalyst is of prime importance in reducing the level of nitrogen-containing organic molecules, which poison the dedicated cracking catalysts that are applied in the second stage. In this second stage full conversion to lighter products is achieved, by applying a liquid recycle. Differences between both modes are described elsewhere [31, 32]. Determination of the catalyst behavior with time on stream under conditions simulating commercial practice, as an alternative to rapid catalyst stability screening with its shortcomings [3, 33, 34], is of prime importance to determine beforehand the achievable cycle length of the hydrocracker. However, the measured performance has to be extrapolated significantly to evaluate commercial end-of-run conditions. This extrapolation should be carried out by using fit for purpose models, which are a trade-off between simplicity (related to the available, or generated, kinetic data) and kinetic detail (related to the required value of the extrapolation). In principle, two types of process models can be distinguished. In correlative models the effect of certain process (or catalyst) variables is incorporated in mathematical formulas, which in themselves may have limited or no physical meaning. They are an

Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 3357

Figure 6. Deactivation of HDN activity of first stage hydrocracking catalysts. Figure 8. Decrease of apparent HDN deactivation energy (E) as function of cumulative amount of oil processed.

Figure 7. Model simulations: (A) Cycle length as function of initial HDN activity k(0) and sensitivity to coke deactivation k(d,0). (B) Cycle length and SOR temperature requirement as function of pressure (P).

abbreviated form of the experimental database, relatively easy to set up and convenient to use, but sometimes difficult to expand [11, 20, 35]. Oil companies, including Shell, have also developed more ambitious (and generally more time-consuming to set up) kinetic models, in which the feedstock is split into certain (groups of) components reacting according to a kinetic scheme [3, 36], ultimately to evaluate entire refining schemes. [3,5] Kinetic models based on networks of elementary steps have been made available by Froment and co-workers [37-40] for model feedstocks. For the purpose of this paper a relatively simple kinetic model is used to illustrate catalyst deactivation phenomena. The model, fitted for simulations of trends in cycle lengths as a function of prime operation and catalyst parameters, belongs to the category of correlative models with a pseudokinetic basis. The deactivation of the HDN activity of a typical Ni/Mo/alumina hydrotreating catalyst under first-stage hydrocracking conditions in pilot plant runs is described by pseudohigh-order deactivation; see Figure 6. The model parameters were fitted by using pilot plant deactivation data gathered over a wide range of combinations of fixed pressure and fixed temperature. For a certain combination of catalyst and feed the model describes the rapid decline encountered in the first phase of the runs, followed by a more steady decline. Figure 6 also illustrates, for one P, T combination, the experimental data and model line. The model can be used to elucidate the trends in cycle length development as a function of catalyst activity (parameter k(0), which is still a function of P and T according to first-order kinetics), sensitivity to coke deactivation (parameter k(d,0)), and operating conditions, such as pressure, initial temperature, and space velocity. Figure 7, with the parameters scaled to that at base case conditions, further illustrates that both increasing the resistance to coke deactivation (lowering

k(d,0)) and increasing initial HDN activity (and thus decreasing the SOR temperature requirement) significantly improve the cycle length. The latter widens the available temperature window (from SOR to design EOR temperature), while at the lower temperature side of the window catalyst deactivates slower. This sets guidelines for further catalyst development, i.e., increasing dispersion and number of active sites (HDN and hydrogenation) and improving porous texture. Obviously, an increase of operation pressure has a large effect on the achievable cycle length, both by decreasing the coke formation and increasing the initial HDN activity. High pressure hydrocrackers do require larger capital investments than low pressure hydrocrackers. Nevertheless, when the feedstocks to be processed contain large amount of coke precursors, high pressure hydrocracking might become imperative. Although this correlative deactivation model is very helpful in providing guidelines for catalyst development, additional steps are required for quantitative extrapolation of pilot plant runs of limited length to real commercial cycle lengths. At the same time, the model will then be upgraded step-by-step in the direction of the more sophisticated ones [3, 5, 37-40], including that of Shell. Four obvious extensions of the pseudokinetic basis are the following: 1. It can account for thermodynamical limitations at higher temperatures (>400-410 °C, depending on the pressure). This results in a lower temperature response of the HDN reaction at higher temperatures [32] and thus in shorter calculated cycle lengths. 2. Extension of the one-key component, total Nconcentration-based (pseudo-) kinetic description to a multicomponent description, with different types of N-compounds (e.g., five- and six-membered N-containing rings) will make the model more robust. More detailed analyses of both feed and products are then required. 3. It can account for the increasing pore diffusion limitation effects with increasing coke deposition in the pores. Pilot plant results suggest that the apparent HDN activation energy indeed decreases with increasing amount of oil processed; see Figure 8. Obviously, further work is required to prove this increase of pore diffusion limitation. 4. Because in commercial practice feedstocks to the hydrocracker often change within an operation cycle, effects of feedstock parameters such as heaviness, aromatics, and organic nitrogen contents on the pseudokinetic parameters of both the main and deactivation reaction should be incorporated. The rate at which the HDN activity of the hydrotreating catalyst declines is only one factor influencing the overall cycle length of the hydrocracker. To illustrate

3358 Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997

List of Symbols

Figure 9. Changes in product properties as function of time on stream. (A) Simplified scheme of hydrocracker. (B) Temperature required for fixed conversion as function of time on stream. (C) Monoaromatics content of product fraction as function of time on stream. (D) Hydrocarbon distribution for SOR and EOR.

some of the other factors, the performance of a modern single-stage hydrocracker processing a heavy feedstock is shown in Figure 9. As a result of deactivation of the catalyst system by coke, the operating temperature required for a certain >370 °C conversion level increases with time on stream at a decreasing rate. As result of applying a modern hydrocracking catalyst system, this operating temperature increase had no significant impact on the middle distillate yield. Nevertheless, with increasing temperature, the aromatics content of the combined naphtha, kerosene, and gas oil fractions increases with time on stream, again at a decreasing rate. Although increasing aromatics content is beneficial for the naphtha fraction, it has a negative effect on the prime combustion properties and specifications of the kerosene (smoke point) and gas oil fraction (cetane number). Hydrocarbon-type distributions are clearly different between start-of-run (SOR) and end-of-run (EOR) as illustrated by mass spectroscopy [41]. With time on stream the hydrogen deficiency of the middle distillate fraction increases, i.e., the net transition from naphthenes and paraffins to aromatics. The specifications of the diesel oil and kerosene fractions could then become important as a limitation of the available cycle length. In this paper we have illustrated that for oil refinery process and catalyst development the maximum catalyst life, or cycle length, is of prime importance. The catalyst cycle length is determined by initial activities for the relevant reactions, their decline rates, changes of both product yields and qualities during deactivation, and (equipment) fouling by byproducts. The objectives for development of more stable catalysts are longer cycle lengths in the refineries, intake of heavier feedstocks, and increase of the severity of operation. Stepwise development of improved catalysts in combination with dedicated performance studies and fit-for-purpose models will lead to the required dedicated catalyst systems. Acknowledgment We thank our colleagues Dr. K. P. de Jong and Dr. K. Vanden Bussche (Shell Research and Technology Centre, Amsterdam) for stimulating and critical discussions on the subject of this paper.

a(t) ) deactivation function k(d) ) deactivation rate constant (L/kg) k(d,0) ) pseudodeactivation frequency factor [L/(kg barm)] k′(d) ) pseudo- (pressure independent) deactivation rate constant [L/(kg barm)] k(t) ) rate constant main reaction [kg/(L h)] k(0) ) initial rate constant main reaction [kg/(L h)] k(HDS) ) hydrodesulfurization rate constant m ) pressure dependency exponent n ) deactivation order t ) time on stream (h) E ) apparent activation energy main reaction (J/mol) E(d) ) pseudodeactivation activation energy (J/mol) EOR ) end of run; end of operation cycle HC ) hydrocracking HDN ) hydrodenitrogenation HDS ) hydrodesulfurization HT ) hydrotreating MD ) middle distillates (kerosene + gas oil) MHC ) mild hydrocracking P ) pressure (bar) SOR ) start of run; start of operation cycle SV ) space velocity [kg/(L h)] T ) temperature (°C) (base) ) for the base case (var) ) variation relative to base case

Literature Cited (1) Butt, J. B.; Petersen, E. E. Activation, Deactivation and Poisoning of Catalysts; Academic Press: New York, 1988. (2) Bartholomew, C. H. In Catalytic Hydroprocessing of Petroleum and Distillates; Oballa, M. C., Shih, S. S., Eds.; Marcel Dekker: New York, 1994. (3) Stork, W. H. J. Deactivation and performance testing of hydroconversion catalysts. Prepr. Am. Chem. Soc., Div. Pet. Chem. 1995, 40, 547 and references cited therein. (4) Fetzer, J. C. The role of the Scholl condensation in the chemistry of catalytic hydrocrackers. Polycyclic Aromat. Compounds 1994, 5, 19. (5) van der Eijk, H.; Vink, K. J. Paper presented at the 8th Int. Symposium on Large Chemical Plants, Antwerp, Belgium, 1992. (6) Gosselink,J. W.; Stork, W. H. J. A simple multicomponent description of the influence of the particle and pore size of commercial catalysts on the hydrodesulfurization of heavy gas oil in (small) trickle-flow reactors. Chem. Eng. Process. 1987, 22, 157. (7) Durante, V. A. Polycyclic aromatic cleavage (parc) process. European patent 0 564 254, 1993. (8) Absi-Halabi, M.; Stanislaus, A. Coke formation on catalysts during the hydroprocessing of heavy oils. Appl. Catal. 1991, 72, 193. (9) de Jong, K. P.; Reinalda, D.; Emeis, C. A. Coke deposition in trickle-bed reactors during heavy oil processingsCatalytic and physical effects. In: Catalyst Deactivation; Delmon, B., Froment, G. F., Eds.; 1994. Stud. Surf. Sci. Catal. 1994, 88, 155. (10) Richardson, S. M.; Nagaishi, H.; Gray, M. R. Initial carbon deposition on a NiMo/Al2O3 bitumen hydrocracking catalystThe effect of reaction time and hydrogen pressure, Prepr.-Am. Chem. Soc., Div. Petr. Chem. Int. Symp. Deactivation Testing Hydrocarbon Conversion Catalysts; Aug 20-25, 1995; pp 455459. (11) Yamamoto, Y.; Kumata, F.; Massoth, F. E. Hydrotreating catalyst deactivation by coke from SRC-II oils. Fuel Process. Technol. 1988, 19, 253. (12) Beyne, A. O. E.; Froment, G. F. A percolation approach for the modeling of deactivation of zeolite catalyst by coke formation. Chem. Eng. Sci. 1990, 45, 2089. (13) Muhammad, O. H. J.; Kam, E. K. T. Modeling and experimental studies of coking in a structured catalyst. Chem. Eng. Sci. 1996, 51, 2939. (14) Guisnet, M.; Magnoux, P. Coking and deactivation of zeolites. Influence of the pore stucture. Appl. Catal. 1989, 54, 1.

Ind. Eng. Chem. Res., Vol. 36, No. 8, 1997 3359 (15) Masuda, T.; Hashimoto, K. Deactivation of Zeolite Catalysts Caused by Coke Deposition. Prepr. Symp. Deactivation and Testing of Hydrocarbon Conversion Catalysts; 1995, 40, 359. (16) George, S. E.; Boardman, S. R.; Foley, R. M.; Sanborn, L. J.; Johnson, P. S.; Webb, A.; Gallagher, A.; Gualtieri, P. K.; Mok, W. S.; Nash, D. Paper AM 94-19, NPRA Meeting, March 20-22, 1994, San Antonio, TX. (17) Hadjiloizou, G. C.; Butt, J. B.; Dranoff, J. S. Piperidine hydrogenolysis on a commercial hydrocracking catalyst. J. Catal. 1992, 135, 481. (18) Latif, N. M. A. Polynuclear aromatic formation in hydrocracker and their impact on catalyst stability. In: Catalysis in Petroleum Refining 1989; Trimm, D. L., et al. Eds.; Elsevier Science Publishers B. V.: Amsterdam, 1990. (19) Gosselink, J. W.; Stork, W. H. J.; de Vries, A. F.; Smit, C. H. Mild hydrocracking: Coping with catalyst deactivation. In: Catalyst Deactivation 1987; Delmon, B., Froment, G. F., Eds.; Elsevier Science Publishers B. V.: Amsterdam, 1987. (20) Gosselink, J. W.; van de Paverd, A.; Stork, W. H. J. Mild hydrocracking: Optimization of multiple catalyst systems for increased vacuum gas oil conversion. In: Catalysis in Petroleum Refining 1989; Trimm, D. L., et al., Eds.; Elsevier Science Publishers B. V.: Amsterdam, 1990. (21) Brown, T. Cracking cooperation, Pte. Rev. V40 N.470 4, March 1988. (22) Hamilton, G. L.; Fitzgerald, M.; Damico, V. Retrofitting heavy oil processes. AIChE 1986 Annu. Meet., Miami Beach 11/ 2-7/86; Prepr. N.7A 52P. (23) Mawandza, N.; Rouquier, R. Converting a hydrotreater to a low pressure hydrocracker. Hydrocarbon Technology International 57-58, 1994. (24) Sebestyen, B.; Deak, A. Environmentally friendly motor fuels in Hungary. 14th World Petroleum Congress; Stavanger, Norway 5/29-6/1/94; preprint N.9.4, 1994. (25) Plantenga, F. L.; Sonnemans, J. W. M. Hydroconversion of vacuum gas oil and atmospheric residua, 185th ACS Natl. Meeting; Seattle 3/20-25/83; ACS Div. Pet. Chem. Prepr. V28 N.3 621-32; March 1983. (26) Stork, W. H. J.; Lundqvist, L.; Schaper, B.; Maxwell, I. E.; de Vries, A. F. Mild hydrocracking, an attractive way of adding conversion. 27th DGMK Annu. Meeting; Aachen 10/6-8/82; Deutsch. Ges. Mineraloelwiss. KohleChem. E. V. Compend. 85-87; 1982-1983. (27) Salazar, J. A.; Gonzalez, J. C.; Rodriguez, E.; Morales, A.; Galiasso, R. E. Hydrocracking adjusted for the production of diesel fuel. Rev. Tec. Intevep. V7 N.2 123-33; July 1987. (28) Du, C.; Huang, Z. Investigation on the acidity of zeolite based mild hydrocracking catalysts. J. Fuel Chem. Technol. Sept 1994, 22 (3), 308. (29) Dufresne, P.; Bigeard, P. H.; Billon, A. Mild hydrocracking, a process between hydrotreatment and hydrocracking (for upgrad-

ing vacuum gas oils to more valuable products over special zeolite catalysts): 195th ACS Natl. Meeting -3rd Chem. Congr. North Am.; Toronto 6/5-10/88; Abstr. Pap. Abstr. No. I&EC-213 V1, 1988. (30) Huang, Z.; Zhang, X.; Li, C.; Shi, J. Selection of zeolitic acid component for mild hydrocracking catalysts for processing poor quality VGO. Pet. Process. 1992, 3, 11. (31) van Dijk, A.; de Vries, A. F.; van Veen, J. A. R.; Stork, W. H. J.; Blauwhoff, P. M. M. Evaluation of hydrocracking catalysts in recycle tests. Catal. Today 1991, 11, 129. (32) Minderhoud, J. K.; van Veen, J. A. R. First-stage hydrocracking: Process and catalytic aspects. Fuel Process. Technol. 1993, 35, 87. (33) Pernicone, N. Methods for laboratory-scale evaluation of catalyst life in industrial plants. Appl. Catal. 1985, 15, 17. (34) Delmon, B. Characterization of catalyst deactivation: Industrial and laboratory time scales. Appl. Catal. 1985, 15, 1. (35) Ternan, M.; Kriz, J. F. Some effects of catalytic composition on deactivation and coke formation when hydrocracking Athabasca bitumen. In: Catalyst Deactivation; Delmon, B., Froment, G. F., Eds.; Elsevier Science Publishers B. V.: Amsterdam, 1980. (36) Quann, R. J.; Jaffe, S. B. Building useful models of complex reaction systems in petroleum refining. Chem. Eng. Sci. 1996, 51, 1615. (37) Baltanas, M. A.; Froment, G. F. Computer generation of reaction networks and calculation of product distributions in the hydroisomerization and hydrocracking of paraffins on Pt-containing bifunctional catalysts. Comput. Chem. Eng. 1985, 9, 71. (38) Froment, G. F. The kinetics of complex catalytic reactions. Chem. Eng. Sci. 1987, 42, 1073. (39) van Raamdonck, K. K.; Froment, G. F. Fundamental kinetic modeling of hydroisomerization and hydrocracking of paraffins on noble metal loaded zeolites. II The elementary cracking steps. AIChE-meeting, San Francisco, Nov 5-10, 1989. (40) Baltanas, M. A.; van Raamsdonck, K. K.; Froment, G. F.; Mohedas, S. R. Fundamental kinetic modeling of hydroisomerization and hydrocracking on noble-metal-loaded faujasites. 1. Rate parameters for hydroisomerization. Ind. Eng. Chem. Res. 1989, 28, 899. (41) Wadsworth, P. A.; Villalanti, D. C. Pinpoint hydrocarbon types. Hydrocarbon Process. Int. Ed. 1992, 71, 109.

Received for review September 30, 1996 Revised manuscript received May 9, 1997 Accepted May 13, 1997X IE9605995

X Abstract published in Advance ACS Abstracts, July 1, 1997.