Ind. Eng. Chem. Res. 2005, 44, 9409-9413
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Kinetic Model for Moderate Hydrocracking of Heavy Oils Sergio Sa´ nchez,† Miguel A. Rodrı´guez,‡ and Jorge Ancheyta*,†,§ Instituto Mexicano del Petro´ leo, Eje Central La´ zaro Ca´ rdenas 152, Me´ xico, D.F. 07730, Facultad de Ingenierı´a, UNAM, Ciudad Universitaria, Me´ xico D.F. 04510, and Escuela Superior de Ingenierı´a Quı´mica e Industrias Extractivas (ESIQIE-IPN), UPALM, Zacatenco, Me´ xico D.F. 07738
In this work, a kinetic model for hydrocracking of heavy oils is proposed. The model includes five lumps: unconverted resid (538 °C+), vacuum gas oil (VGO; 343-538 °C), distillates (204343 °C), naphtha (IBP-204 °C), and gases. Kinetic parameters were estimated from the experimental results obtained in a fixed-bed downflow reactor. The proposed lump kinetic model fitted the data for hydrocracking of Maya heavy crude with a Ni/Mo catalyst at 380-420 °C reaction temperature, 0.33-1.5 h-1 liquid hourly space velocity, 5000 scf/bbl, and 70 kg/cm2. Only resid cracked to naphtha at 380 °C, but all heavier fractions produced some naphtha at higher temperature. The predicted product composition is in good agreement with experimental values with an average absolute error of less than 5%. 1. Introduction Hydrocracking is a process commonly used in the petroleum refining industry to treat heavy oils, such as vacuum resid. The main objective of this process is to convert heavy molecules into lighter and more valuable products. Various technologies are available for the upgrading of heavy oils. Some processes include sufficient large-scale commercial experience to be regarded as well established, mature processes. There are also processes that are in the early stages of commercial application, together with others that have been well tested at the demonstration scale.1 Hydrocracking processes are strongly influenced by the method of feed introduction, the arrangement of the catalyst beds, and the mode of operation of the reactors. Hence, a proper selection and design of reactors are very important for this. Depending on the nature of the feed, the reaction is generally carried out in fixed-bed, moving-bed, or ebullated-bed reactors. Sometimes a combination of different reactors is preferred.2 The most-known hydrocracking technology is the H-oil process, which is capable of maximizing the throughput of heavy feedstocks to produce environmentally friendly lighter products.3 The H-oil process is designed to operate at high severity in an ebullatedbed reactor to produce maximum distillates and minimum fuel oils.4 The common operating conditions of the H-oil process are 170-200 kg/cm2 pressure, 400-440 °C temperature, and 0.1-0.3 h-1 liquid hourly space velocity (LHSV). At these conditions, the conversion level is higher than 50%.5 The severe hydrocracking conditions cause problems of coke deposition on the catalyst and sludge formation in the product oil, which eventually shorten the life of the catalyst, plug the transfer lines, and deteriorate the quality of the products.6 To avoid the problems associ* To whom correspondence should be addressed at Instituto Mexicano del Petro´leo. Fax: (+52-55) 9175-8429. E-mail:
[email protected]. † Instituto Mexicano del Petro´leo. ‡ UNAM, Ciudad Universitaria. § Escuela Superior de Ingenierı´a Quı´mica e Industrias Extractivas (ESIQIE-IPN), UPALM.
ated with sediment formation, The Mexican Institute of Petroleum (IMP, Instituto Mexicano del Petroleo) has developed a process for the hydrotreating of heavy petroleum oils, which, among several characteristics, operates at moderate reaction conditions and improves the quality of the feed, keeping the conversion level low.7,8 These conditions change the selectivity of all of the reactions, especially for those of hydrocracking. There are various kinetic models for the hydrocracking of heavy oils reported in the literature and these were recently summarized elsewhere.9 Reported models are commonly developed for severe reaction conditions. Only one model10 makes the differentiation of different hydrocracking regimes, namely, low, intermediate, and high hydrocracking, but it does not take into account gas formation. Different approaches have been utilized for kinetic modeling of hydrocracking, varying from the most common and used lumping technique to more complex models based on continuous mixtures11 or single events.12,13 Govindakanaan and Froment14 also applied the single-event kinetic modeling to the hydrocracking of vacuum gas oil (VGO) on a Pt zeolite. The authors recognized that the equations become somewhat more complex because of the multiphase operation of hydrocrackers. In the case of lumped kinetics, models with only a few lumps have been proposed in the literature (mainly resid, distillates, and naphtha) and other important lumps, such as VGO (common feed to catalytic cracking units) and gases, are not common in hydrocracking lumped models. To have a better understanding of the IMP process for the hydrotreating of heavy oils, in this work we conducted some experiments in a pilot plant equipped with a downflow fixed-bed reactor, with the main objective of obtaining information for developing a kinetic model for the hydrocracking of heavy oils at moderate reaction conditions. The proposed model, which is based on the yet acceptable lumping approach, can be applied for catalyst screening and basic process studies. A more complex and detailed kinetic model is under development and will be presented in further papers.
10.1021/ie050202+ CCC: $30.25 © 2005 American Chemical Society Published on Web 06/15/2005
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Figure 1. Experimental setup for moderate hydrotreating experiments.
2. Experimental Section 2.1. Brief Description of the Experimental Setup. All of the experiments were carried out in a fixed-bed high-pressure pilot plant, which is shown schematically in Figure 1. Hydrotreating was conducted in oncethrough hydrogen in a downflow mode of operation. A detailed description of the pilot plant, reactor, catalyst loading, experimental procedure, and catalyst activation was reported elsewhere.15,16 The heart of the pilot plant is an isothermal reactor (an inside diameter of 2.54 cm and a total length of 143 cm). The temperature of the reactor is maintained at the desired level by using a three-zone electric furnace, which provided an isothermal temperature along the active reactor section. The temperature profile is measured by a movable axial thermocouple located inside the reactor. 2.2. Catalyst Loading and Activation. A Ni/Mo commercial catalyst was employed for all experiments (175 m2/g specific surface area, 0.56 cm3/g pore volume, 127 Å mean pore diameter). The catalyst was first crushed and sieved (100 mesh) in order to minimize interphase mass-transfer resistances. A total of 100 mL of catalyst was loaded into the reactor and in situ activated by sulfiding with a hydrodesulfurized naphtha containing 0.8 wt % CS2 at the following operating conditions: pressure of 54 kg/cm2, H2/oil ratio of 2000 ft3/bbl, temperature of 230 °C, and LHSV of 3.2 h-1. The sulfiding time was 18 h. 2.3. Hydrotreating Experiments and Product Analysis. Maya heavy crude oil was employed as the feed and has the following main properties: 21.97° API gravity and 3.51 wt % sulfur, 10.5 wt % Ramsbottom carbon, 12.4 wt % n-heptane-insoluble asphaltenes, 292
wppm V, and 53 wppm Ni contents. The weight percentages of distillates of this crude are 12.1 wt % naphtha, 20.4 wt % distillates, 28.1 wt % VGO, and 39.4 wt % resid. The experiments were carried out at the following conditions: 380-420 °C temperature and 0.33-1.5 h-1 LHSV, keeping constant the hydrogen-tooil ratio and reactor pressure at 5000 ft3/bbl and 70 kg/ cm2, respectively. The products were analyzed by simulated distillation following the ASTM D-5307 method. The hydrocracked products were defined as follows: gases, naphtha (IBP; 204 °C), distillates (204-343 °C), VGO (343-538 °C), and unconverted resid (538 °C+). 3. Results and Discussion 3.1. Variation in the Product Composition. Figure 2 is an example of the variation of reaction conditions on distillation curves. Maya heavy crude and Isthmus light crude distillation curves are also shown for comparison. It is clearly observed that, as the LHSV is reduced and consequently the reaction severity is increased, the distillation curves are moved to the right, which means that high boiling point molecules are converted into lighter ones. For the data shown in Figure 2, the total amounts of distillates in weight percent recovered up to 538 °C are 60.5% for Maya crude, 62.6% for LHSV ) 1.25 h-1, 64.2% for LHSV ) 1.0 h-1, 65.7% for LHSV ) 0.75 h-1, and 78.3% for Isthmus crude. This variation in distillation curves also changes the conversion and product composition. For all of our experiments, the conversion was kept below 50%, and
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Figure 5. Proposed kinetic model for the hydrocracking of heavy oils. Figure 2. Effect of the LHSV on the distillation curve of hydrotreated products at 380 °C: (s) Maya; (‚‚‚) Isthmus; (O) LHSV ) 1.25 h-1; (b) LHSV ) 1.0 h-1; (0) LHSV ) 0.75 h-1.
Figure 3. Effect of the LHSV and temperature on the conversion: (O) 380 °C; (b) 400 °C.
Figure 4. Product composition as a function of the LHSV at 400 °C: (O) VGO; (b) distillates; (0) naphtha; (9) gases.
consequently sediment formation in all of the products was lower than 0.05 wt %. The variation of the conversion as a function of the temperature and LHSV is presented in Figure 3. This conversion is defined as follows:
conversion ) 538 °C+ in feed - 538 °C+ in product × 100 (1) 538 °C+ in feed The product composition, determined as the grams of each product divided by the total grams of the reactor effluent except hydrogen, is shown in Figure 4. The expected behavior is observed; that is, the higher the temperature and the lower the space velocity, the higher the conversion and product yields.
It is also seen from Figures 3 and 4 that the effect of LHSV on the conversion and product yields at values lower than 1.0 is minimal, which implies that the equilibrium conversion is almost reached. This also means that, under moderate hydrocracking, the regime reaction severity must not be higher than that employed in our experiments. If the severity is increased, the conversion will achieve values higher than 50% and consequently sediment formation will also be high at levels where commercial plant operation has to be stopped. Another observation from these figures is that the reaction selectivity is mainly oriented toward the production of VGO and distillates. Naphtha production is only slightly higher than that of the original feed. This behavior has two explanations: (1) the naphtha formation rate is almost equal to the naphtha hydrocracking rate, or (2) naphtha formation from heavy fractions is insignificant. By this means, i.e., only observing the experimental results, it is not possible to establish which of the two mechanisms is prevailing. However, calculation of the kinetic parameters for each reaction path could help in this task, which is explained in the following section. 3.2. Kinetic Modeling. Figure 5 shows the proposed kinetic model, which includes 5 lumps (unconverted resid, VGO, distillates, naphtha, and gases) and 10 kinetic parameters (k1, ..., k10). For each reaction, a kinetic expression (ri) was formulated as a function of the product composition (yi) and kinetic constants (ki). Product compositions were determined with pilot-plant mass balances and distillation curves. All reactions were assumed to be first order. On the basis of these considerations, the reaction rates of the proposed model are:
Resid: rR ) -(k1 + k2 + k3 + k4)yR
(2)
VGO: rVGO ) k1yR - (k5 + k6 + k7)yVGO
(3)
Distillates: rD ) k2yR + k5yVGO - (k8 + k9)yD
(4)
Naphtha: rN ) k3yR + k6yVGO + k8yD - k10yN
(5)
Gases: rG ) k4yR + k7yVGO + k9yD + k10yN
(6)
The kinetic model was incorporated into an isothermal plug-flow reactor model. On the basis of previous experiences, axial dispersion and external and internal gradients are neglected.17 The following mass balance, solved with a Runge-Kutta method, was used to eval-
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Table 1. Kinetic Parameters of the Proposed Model temperature 400 °C 420 °C
kinetic constant (h-1)
380 °C
k1 k2 k3 k4
0.042 0.008 0.008 0.041
Resid 0.147 0.022 0.020 0.098
0.362 0.057 0.043 0.137
48.5 44.2 38.0 27.3
k5 k6 k7
0.018 0 0
VGO 0.057 0.104 0.007 0.016 0 0
39.5 37.1
k8 k9
0 0
Distillate 0.003 0.010 0 0
53.7
0
Naphtha 0 0
k10
activation energy EA (kcal/mol)
Figure 6. Arrhenius plot for the different kinetic parameters: (O) k1; (b) k2; (0) k3; (9) k4; (4) k5; (2) k6; (]) k8.
uate the product composition from a set of kinetic constants for each temperature:18
dyi d(1/LHSV)
) ri
(7)
The minimization of the objective function, based on the sum of square errors between the experimental and calculated product compositions, was applied to find the best set of kinetic parameters. This objective function was solved using the least-squares criterion with a nonlinear regression procedure based on Marquardt’s algorithm.19 Table 1 summarizes the values of the kinetic parameters together with the activation energies for each reaction. The first finding from the results presented in Table 1 is the confirmation of what we were discussing in the last paragraph of the previous section: naphtha hydrocracking is insignificant at the conditions of this work because the values found for k10 are null. This means that naphtha formation comes from the hydrocracking of heavy fractions, especially from the resid fraction, but at a very small rate. Given that k4 * 0 and k7 ) k9 ) k10 ) 0, it is then concluded that gas production is exclusively from the resid at the temperature range of this study, 380-420 °C. Hydrocracking selectivity slightly changes at the different temperatures. For instance, at 380 °C no formation of naphtha from VGO and distillates is observed because the values of k6 and k8 are zero. On the contrary, at 400 and 420 °C the values of these parameters are different from zero. Because some values of the kinetic parameters were found to be zero, not all of the activation energies could be estimated. Figure 6 shows the Arrhenius plot for all kinetic constants. The correlation coefficients for all cases were virtually unity. The values of the activation energies for some reactions are also presented in Table 1 and are within the range of those reported in the literature.9,10 A comparison of the experimental product composition and those determined by solving eqs 2-7 with the ki values given in Table 1 is shown in Figure 7. It is observed that the product composition is quite well predicted with an average absolute error of less than 5%, which indicates that the proposed kinetic model is adequate for the moderate hydrocracking of heavy oils. From the results presented in this investigation, it should be highlighted that hydrocracking can be carried
Figure 7. Comparison between the experimental and calculated product compositions: (O) resid; (b) VGO; (0) distillates; (9) naphtha; (4) gases.
out at moderate operating conditions in order to avoid sludge and sediment formation by keeping the conversion below than 50%. At these conditions, the hydrocracking reaction is not deep enough to increase the yields of light products (naphtha and gases) and only the production of distillates and VGO is affected. The decision of operating or designing the commercial hydrocracking units at the proposed reaction conditions will be dictated by further detailed technical and economical studies. 4. Conclusion A five-lump kinetic model for the moderate hydrocracking of heavy oils was developed. The proposed model is capable of predicting the production of unconverted resid, VGO, distillates, naphtha, and gases, with an average absolute error of less than 5%. Acknowledgment The authors thank IMP for financial support. M.A.R. also thanks CONACyT for financial support. Nomenclature rR ) reaction rate of resid (wt %/h) rN ) reaction rate of naphtha (wt %/h) rD ) reaction rate of distillates (wt %/h) rVGO ) reaction rate of VGO (wt %/h) rG ) reaction rate of gases (wt %/h) k1 ) first-order rate constant for the hydrocracking of resid to VGO (h-1)
Ind. Eng. Chem. Res., Vol. 44, No. 25, 2005 9413 k2 ) first-order rate constant for the hydrocracking of resid to distillates (h-1) k3 ) first-order rate constant for the hydrocracking of resid to naphtha (h-1) k4 ) first-order rate constant for the hydrocracking of resid to gases (h-1) k5 ) first-order rate constant for the hydrocracking of VGO to distillates (h-1) k6 ) first-order rate constant for the hydrocracking of VGO to naphtha (h-1) k7 ) first-order rate constant for the hydrocracking of VGO to gases (h-1) k8 ) first-order rate constant for the hydrocracking of distillates to naphtha (h-1) k9 ) first-order rate constant for the hydrocracking of distillates to gases (h-1) k10 ) first-order rate constant for the hydrocracking of naphtha to gases (h-1) yR ) resid composition (wt %) yN ) naphtha composition (wt %) yD ) distillates composition (wt %) yVGO ) VGO composition (wt %) yG ) gases composition (wt %)
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Received for review February 19, 2005 Revised manuscript received April 28, 2005 Accepted April 29, 2005 IE050202+
