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Hydrocracking of vacuum residue with solid and dispersed phase catalyst: Modeling of sediment formation and hydrodesulfurization
/ Fuel Processing Technology 159 (2017) 320–327
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Fuel Processing Technology journal homepage: www.elsevier.com/locate/fuproc
Research article
Hydrocracking of vacuum residue with solid and dispersed phase catalyst: Modeling of sediment formation and hydrodesulfurization Eduard Manek, Juma Haydary ⁎ Institute of Chemical and Environmental Engineering, Faculty of Chemical and Food Technology, Slovak University of Technology, Radlinského 9, 812 37 Bratislava, Slovakia
a r t i c l e
i n f o
Article history: Received 13 December 2016 Received in revised form 31 January 2017 Accepted 5 February 2017 Available online xxxx Keywords: Residual hydrocracking Sediment formation Hydrodesulfurization Mathematical modeling
a b s t r a c t Effect of dispersed Mo based catalyst on residual hydrocracking was investigated in this paper. The study was performed in an industrial scale hydrocracking unit with the average reactor temperature of 399–419 °C and pressure of 18 MPa. Vacuum residue from Ural crude oil was fed to the reactor. A sediment formation model was proposed with parameters sets for the regime with classical supported Ni-Mo/Al2O3 catalyst and for that with a combined system of supported and dispersed catalyst. Four process parameters were incorporated in the model equation: reaction temperature, hydrogen partial pressure, exothermic gain and hydrogen makeup to feed mass ratio. A decrease of the sediment formation with the application of dispersed Mo catalyst was observed, which results in higher Mo concentration in the liquid phase and thus in higher hydrogenation activity and stabilization of coke precursors. Hydrodesulfurization was also investigated and a sulfur distribution model was created with parameters for both regimes. The model contains three lumps with significant sulfur content: sulfur contained in the vacuum residue, sulfur contained in the vacuum gas oil and hydrogen sulfide. Application of the catalyst and increase of reaction temperature from 399–414 °C to 414–419 °C resulted in increase of average sulfur conversion in vacuum residue from 86.5% to 87.9%. © 2017 Published by Elsevier B.V.
1. Introduction Hydrocracking is a well-established refining process for upgrading heavy feedstock into valuable products such as gasoline, kerosene and diesel oil. First commercial hydrocracking units operation began during the interwar period [1], many enhancements have been proposed since then. One of the current challenges is to utilize the growing share of heavy, extra-heavy and alternative feedstocks due to the depletion of light petroleum reserves [2,3]. These materials contain high amounts of metals (Ni, V), organo-sulfur compounds and asphaltenes which cause operational problems. Classical solid hydrocracking catalysts cannot effectively process these impurities, and the application of dispersed phase catalysts draws more attention. Several studies and reviews dealing with slurry or dual phase catalysis (the use of both supported and dispersed catalysts) of heavy residue hydroconversion have been published in the past few years [4–9]. Active component of these dispersed phase catalysts is a metal sulfide (mostly MoS2) generated in situ by thermal decomposition of the precursor (metallic oxides, organo-metallic compounds). Panariti et al. studied the performance of different oil-soluble precursors and introduced the following order of activity Mo N Ni ~ Ru N Co N V N Fe. Advantages of dispersed catalysts over supported ones is their lower particle size, better mixing in the reaction ⁎ Corresponding author. E-mail address: [email protected] (J. Haydary).
http://dx.doi.org/10.1016/j.fuproc.2017.02.003 0378-3820/© 2017 Published by Elsevier B.V.
system, reduced hot spots and concentration gradient, promotion of hydrogen transfer, which also stabilizes the mixture and suppresses excessive cracking and coke formation [4]. In combination with an acid catalyst, the presence of dispersed MoS2 protects the main cracking catalyst long enough thus allowing the exploitation of unconventional resources [5]. The main limitation in the commercialization of this application is the high investment and operating cost caused by catalyst makeup or complicated catalyst recovery. Bacaud, in his review [9], summarized the expected properties and criteria for a convenient dispersed catalyst as follows: hydrogenation catalyst, active phase is stable under H2S/H2 pressure, high intrinsic activity, low propensity to agglomeration, sintering and sedimentation, cheap and easily disposable. Coke and sediment formation as one of the main bottlenecks of residual oil hydrocracking has drawn the attention of both academic and corporate researchers. Carbonaceous deposits can deactivate the catalyst by poisoning of the active sites or by pore blockage [10]. Retention of carbonaceous molecules on the catalysts is mainly due to their strong adsorption and low volatility (gas-phase reactions) or low solubility (liquid-phase reactions). At high temperatures (N 350 °C), the coke components are mainly polyaromatics. Their formation involves also hydrogen transfer (acid catalysts) and dehydrogenation (bifunctional catalysts) steps in addition to the condensation and rearrangement steps. On microporous catalysts, the retention of coke molecules is caused by their steric blockage within the micropores [11]. Tailleur et al. [12] studied pore plugging effects in a hydrocracking catalyst in an
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ebullated bed reactor. The obtained results indicated that the initial deactivation is mainly due to coke depositions, while its impact on the mass transfer reaction control depends on temperature. In long-term deactivation, metal deposition plays a more important role in blocking the internal micro- and meso-structures and in the formation of an external layer of pellets. Besides catalyst deactivation, coke sedimentation, which occurs in all heavy oil processing equipment, leads to other various problems [13–16]. Fouling in a heat exchanger network reduce the amount of heat recovered from product streams, which have to be compensated by higher furnace fuel consumption and subsequently higher deterioration of the furnace equipment. As the maximum heating capacity or hydraulic limit (due to pipe plugging) in the heat exchanger network is reached, throughput of the whole unit is reduced or in some cases the unit has to be shut down. In the highly competitive environment of petroleum industry, unplanned shutdowns have unfavorable consequences. High fouling rates lead to increased equipment cleaning and cost. The most common mechanism of carbonaceous deposits is free radical polymerization, where unsaturated hydrocarbons react to form longer chain molecules. This mechanism can be inhibited by hydrogenation conditions or by polymerization inhibitors stabilizing the free radicals and terminating propagation reactions [14]. Condensation reactions and asphaltene precipitation also contribute to the sediment deposition. Stanislaus et al. [15] analyzed sediments produced during VR hydrocracking in a pilot plant test. The analyzed sediments contained mainly organic materials, large part of it being sufficiently volatile to boil in the same range as heavy gas oil. A large portion (about 70 wt%) of the sediment is soluble in various solvents. Common practice is to continually apply a dispersant or solvents to prevent particle agglomeration and the formation of large clusters which are deposited more easily. Storm et al. [16] observed sediment formation during laboratory hydrotreating experiments of vacuum residue. Correlations were found between the amount of sediment produced per unit of weight of the vacuum residue feed and four chemical characteristics of the vacuum residue: degree of condensed polynuclear aromaticity, average number of alkyl-groups substituting the polynuclear aromatics, ratio of heptane insolubles to pentane insoluble-heptane solubles, and H/C ratio of the latter fraction. Majority of papers dealing with heavy oil hydrocracking provides experimental data from laboratory scale reactor or pilot plants. Modeling based on industrial scale units suffers from obvious complications and limitations. Process parameters like the reactor temperature are limited by coke generation and catalyst deactivation. Step tests represent an economical burden. Measuring instruments are prone to decalibration during a long run. Composition of the feed may vary over time due to the composition of the crude oil supplied or the performance of adjacent distillation units. Large volumes of the pipeline, columns, reactor, and separator vessels cause long delays between the changes of process parameters and the measured output. For large continuous processes it is also much more difficult to determine the exact reaction time than for small batch processes. Kinetic models were applied to optimize industrial scale hydrocracker units [17–20]. Sildir et al. [17] developed a continuous non-isothermal model for an industrial hydrocracker unit with the ability of real time optimization and online model parameter tuning. The reactor model was later extended with an additional fractionator model and a cascade model predictive control structure was developed to operate both the reactor and the fractionation column at maximum profits [18]. Zhou et al. [20] used ASPEN PLUS to predict feedstock properties and applied a lumped model to determine the optimal reactor bed temperatures for maximal diesel and kerosene yields. Such model also incorporates simplified HDN and HDS kinetics with nitrogen and sulfur as single lumps. In a previous paper [21], a lumped model of an industrial scale residual hydrocracker unit was developed. The model was implemented in the ASPEN PLUS environment to simulate the fractionation section. Simulated product distillation curves showed good
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agreement with the results of plant analysis. Main objectives of the present work are: – find the correlation between sediment formation and process parameters – extend the existing reaction network with a sediment lump and to estimate the model parameters – compare sediment formation for regimes with and without dispersed catalyst – develop a multi-lump HDS model and compare the HDS performance for regimes with and without dispersed catalyst Nomenclature [] mass fraction HDS hydrodesulfurization RHC residual hydrocracker LCO light cycle oil MCB main column bottom FCC fluid catalytic cracker sed sediments preexponential factor, min−1 Ai Eai activation energy, kJ mol−1 G offgas GLN naphtha GO gas oil Ke kerosene kinetic constant, min−1 ki R gas constant, J mol−1 K−1 t reaction (retention) time, min T temperature, °C VGO vacuum gas oil VR vacuum residue w mass fraction n9, n10, nH2, nExo, nφ model parameters P overall pressure, kPa hydrogen partial pressure, kPa PH2 exothermic temperature gain, °C Texo first stage average bed temperature, °C TR1 Ф hydrogen makeup to feed mass ratio sulfur content in vacuum residue (unc - unconverted) SVR sulfur content in vacuum gas oil SVGO 2. Experimental data All data in this work were collected from an industrial scale residual hydrocracker unit (RHC). Process parameters (flow, temperature, pressure) were measured by standard industrial measuring instruments and physicochemical (sulfur content - STN EN ISO 8754, sediments - ASTM D4870) properties were analyzed by laboratory methods. The feed for the residual hydrocracker unit is the vacuum residue from primary distillation of Ural crude oil. Vacuum residue is the heaviest fraction obtained from primary crude oil distillation, which contains substances with high molecular weight (about 600 × kg kmol−1), its distillation curve is in the range of 440 + °C, and the density is about 1020 kg m− 3 at 20 °C. High sulfur content (about 3.00 mass%), nitrogen (0.45 mass%), and metals (vanadium: 210 mg kg−1, nickel: 66.5 mg kg−1, and sodium: 4.35 mg kg−1). The feed is mixed with heavy aromatic oils, light cycle oil (LCO) and the main column bottom (MCB) from the FCC unit to suppress coagulation of asphaltenes, sediment formation and to maintain required hydrodynamic characteristics of the feed. Diluents represent almost 12 mass% of the whole feed mixture. Before entering the reactor cascade, preheated feed is mixed with high temperature hydrogen. The industrial reactor cascade (Fig. 1) consisted of three identical ebullated bed reactors (diameter of 3.6 m, height of 36.5 m, ebullated bed of approximately 28 m) with a commercial Ni-Mo/Al2O3 hydrocracking catalyst (pellets with a 1 mm diameter and 3.2 mm length).
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Fig. 1. Cascade of ebullated bed reactors.
Temperature in the reactors was held at around 401–419 °C and the pressure at 18–20 MPa. Effluents from the first to the second stage and from the second to the third stage were quenched by hydrocarbon oil and hydrogen to maintain the temperature of the ebullated beds constant and also to compensate for spent hydrogen. Hydrogen makeup is controlled based on the pressure set point in the 1st stage suction drum of a four stage compressor. In the 2nd, 3rd and 4th stage, both hydrogen make-up and recycle gas from the gas-liquid separator are compressed. The process includes on-stream catalyst addition and withdrawal avoiding the reactor shut down for catalyst replacement, which is critical because the catalyst bed is continuously deactivated by metal adsorption and coke deposits. The ebullated bed reactor system is comprehensively described by Martínez et al. [22]. The product stream from the third reactor enters a high pressure separator, where liquid and gas phase are separated. Both streams are subsequently cooled, depressurized and separated into final product fractions in an atmospheric and a vacuum column. Except for the classical solid Ni-Mo/Al2O3 catalyst, an additional commercial Mo-based dispersed phase catalyst was tested with solid catalyst present in the ebullated bed. Oil soluble precursor were injected to the LCO stream at a constant rate, overall mass content of catalyst in the feed was 0.021 mass%. The main purpose of this catalyst is to promote hydrogenation, asphaltene stabilization, to upgrade of the product quality, lower the downstream equipment fouling and to enable operation at higher reactor temperatures without excessive coke formation. Sediments were analyzed in the unconverted residue via an ASTMD4870. Analyzed sediments consisted of both organic and inorganic types. The inorganic part consisted of rust, sand, and crushed catalyst particles; the organic sediments are formed of coke and asphaltenes. Hydrotreating affected only the formation of the organic part; it can be assumed that only a negligible amount of inorganic sediments formed. Sulfur content in both feed and products was analyzed. A significant amount of sulfur was present in the vacuum residue feed, unconverted vacuum residue and in the vacuum gas oil. Sulfur content in other products was in ppm and it was not included in the mathematical model.
vacuum residue and vacuum gas oil mass fraction. d½VR ¼ −ðk1 þ k2 þ k3 þ k4 þ k7 Þ½VR dt
ð1Þ
d½VGO ¼ k4 ½VR−ðk5 þ k6 þ k8 Þ½VGO dt
ð2Þ
d½GO ¼ k3 ½VR þ k5 ½VGO dt
ð3Þ
d½Ke ¼ k7 ½VR þ k8 ½VGO dt
ð4Þ
d½GLN ¼ k2 ½VR þ k6 ½VGO dt
ð5Þ
d½G ¼ k1 ½VR dt
ð6Þ
3.1. Sediment formation model
P H2 nH2 T Exo nExo nφ d½sed : :Ф ¼ k9 :½VRn9 þ k10 :½VGOn10 : P T R1 dt
ð7Þ
Additional process parameters were included in Eq. (7) for the determination of the sediment formation rate which is based on the mass fraction of the vacuum residue ([VR]) and the vacuum gas oil ([VGO]). Initial value of the sediment mass fraction ([sed]) of 0 and irreversible reaction were assumed. Hydrogen partial pressure (PH2) was
3. Mathematical modeling The lumped model of vacuum residue hydrocracking proposed in our previous work consists of six fraction lumps: vacuum residue - VR, vacuum gas oil - VGO, gas oil - GO, kerosene - Ke, gasoline - GLN, gas – G, and of eight reaction steps (Eqs. (1)–(6)). The model was upgraded to predict sediment formation. It was assumed that sources of organic sediments are the vacuum residue and vacuum gas oil lumps. The upgraded reaction network is shown in Fig. 2. Pathways k9 and k10 were not included in Eqs. (1) and (2) due to insignificant effect of
Fig. 2. Reaction network of vacuum residue hydrocracking [21] with two additional steps of sediment formation.
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measured based on the hydrogen concentration and the overall pressure (P) in the gas/liquid separator after the third reactor stage. Partial pressure represents the remaining hydrogen concentration in the gas phase at the end of the reactor cascade. Higher amount of hydrogen should suppress the sediment formation and asphaltene stabilization. Exothermic temperature gain, TExo, across the ebullated bed in the first reactor stage indicated the intensity of hydrogenation; this value is the highest due to the initial hydrogenation reactions in first stage and therefore it was chosen as an equation parameter. To provide dimensionless values, TExo figures in Eq. (7) as ration with average temperature in first stage ebullated bed. Symbol Ф represents the hydrogen makeup to the feed mass ratio before the first reactor stage. Kinetic constants k9 and k10 for sediment formation from the vacuum residue and the vacuum gas oil, respectively, were calculated using the Arrhenius Eq. (8): Eai
ki ¼ Ai :e− RT
ð8Þ
Index i represents the respective reaction pathway. Nine model parameters: A9, A10, Ea9, Ea10, n9, n10, nH2, nExo, nφ, were estimated. Two sets of parameters for both regimes with and without an addition of dispersed phase catalyst were estimated. Objective function was the sum of least squares of real and calculated sediment values. 3.2. Sulfur distribution model In our previous work [21], a simplified sulfur conversion model was proposed. Although this model provides a satisfactory agreement between the calculated and real sulfur conversion, the new model was designed to calculate sulfur distribution considering hydrogen sulfide and the remaining amount in the vacuum residue and the vacuum gas oil fraction, other products contain only an insignificant amount of sulfur. A scheme of the reaction path is shown in Fig. 3. SVR, SVGO and SH2S represents sulfur mass content in vacuum residue, mass content in vacuum gas oil and sulfur mass in hydrogen sulfide. Step ks1 represents the transfer of vacuum residue sulfur to the vacuum gas oil sulfur by an irreversible cracking mechanism. Reversible hydrodesulfurization of the vacuum residue and vacuum gas oil is represented by steps ks1, ks2 −, ks2 +, ks3 −, ks3+. Activation energies and preexponential factors of the Arrhenius Eq. (8) were estimated by minimizing the sum of least squares of real and calculated sulfur content in the vacuum residue and the vacuum gas oil. The amount of sulfur in form of hydrogen sulfide was calculated using the mass balance [Eq. (9)]. Residence time was approximately calculated from the volumetric flow of the feed and the ebullated bed volume. unc Sfeed VR ¼ SVR þ SVGO þ SH2S
Fig. 4. Formation of sediments without dispersed Mo catalyst.
4. Results and discussion 4.1. Analysis of sediment formation Formation of sediments measured during the vacuum residue hydrocracking without the application of dispersed catalyst for over 130 day period is shown in Fig. 4. A 200 day process employing the dispersed catalyst is shown in Fig. 5. A moving average was applied to both calculated and experimental sediment values to eliminate scattering and highlight their trend. The average sediment content in the unconverted vacuum residue without dispersed phase catalyst was 0.28 mass% and the median value was 0.26 mass%, operated in lower severity regime with the average bed temperature of 400–413 °C. The addition of the dispersed phase catalyst decreased the average and the median sediment fraction to 0.24 mass% even in higher severity regime with the average bed temperature of 413–419 °C. A few discrepancies between the real and the predicted sediment values were observed. Inconsistent peaks of experimental and calculated values were visible at days 17, 62, 93, 132, 151 and 180 (Fig. 4) and at days 16 and 127 (Fig. 5), probably due to delayed sediment accumulation and purging from the equipment and pipeline surface, changes of the feedstock quality and other unregistered process deviations. The sediment drop at day 75 for the supported catalyst only regime (Fig. 4), correlates with the rapid decrease of the reactor temperature from 410 °C to 400 °C. Re-establishing the temperature of 412 °C led to an immediate increase of the sediment content. The sediment peak at day 105 for the regime with dispersed catalyst correlates with the increase of the hydrogen partial
ð9Þ
Fig. 3. Reaction scheme for hydrodesulfurization: vacuum residue sulfur content (SVR) and vacuum gas oil sulfur content (SVGO).
Fig. 5. Formation of sediments with dispersed Mo catalyst.
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E. Manek, J. Haydary / Fuel Processing Technology 159 (2017) 320–327 Table 1 Estimated parameters for Eq. (7).
A9 (min−1) Ea9 (kJ mol−1) A10 (min−1) Ea10 (kJ mol−1) n9 n10 nH2 nExo nφ
Supported catalyst
Supported + dispersed catalyst
3.92E+17 242.8 1.05E+23 141.8 476.4 70.52 2.4507 −1.758 6.817
5135 0 1.05E+23 141.8 152.4 175.1 1.456 −1.543 0.9998
4.2. Effect of process parameters on sediment formation
Fig. 6. Parity plot for measured and calculated sediment data.
pressure peak. It is obvious, that the model parameters were estimated based on the degree of change of individual process parameters and the subsequent change of the sediment fraction. Model parameters can be tuned after every non-standard process parameter change to keep the model updated and include new observed effects. A comparison of measured and calculated sediments by a parity plot is given in Fig. 6. Residual plot of measured experimental data is presented in Fig. 7. Regression coefficient for supported catalyst only and combined regime (supported + dispersed) is 0.96 and 0.98, respectively. Residuals are distributed symmetrically in both regimes without any clear pattern. Estimated model parameters for both regimes are shown in Table 1. Due to the very low content of sediments (0.1–0.5 mass%) and the parametric form of Eq. (7), kinetic parameters cannot be compared with values presented in literature. Parameter nH2 for partial hydrogen pressure is almost zero for the regime without dispersed catalyst, which means that the partial hydrogen pressure does not affect the sediment calculation for this regime. Model parameters in this work are plant specific (specific equipment setup, catalyst and feed properties); however, the trend introduced in this study should describe residual hydrocracking in general.
The effects of average bed temperature, hydrogen partial pressure, exothermic temperature gain in the first reactor stage and the hydrogen makeup to feed mass ratio on the sediment were calculated using the presented model. One process parameter was varied while other parameters were set constant at median values. The sediment fraction rapidly increased with the increasing reaction temperature (Fig. 8). Higher temperature intensifies the cracking reactions as well as the radical mechanism of coke formation. Model of the regime employing both supported and dispersed phase catalysts did not show this trend probably due to the promotion of hydrogenation reactions by the dispersed phase catalyst which suppress radical polymerization and stabilize the feedstock. Partial hydrogen pressure in the separator after the third reactor stage showed a significant effect on the sediment formation (Fig. 9) for the regime with both supported and dispersed catalysts. For ideal operation, partial pressure is related to hydrogen make-up. During
Fig. 8. Calculated effect of average bed temperature on sediment formation.
Fig. 7. Residual plot for measured sediment data.
Fig. 9. Calculated effect of hydrogen partial pressure after third stage on sediment formation.
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Table 2 Estimated kinetic parameters for the HDS model.
A1 (min−1) Ea1 (kJ mol−1) A2+ (min−1) Ea2+ (kJ mol−1) A2− (min−1) Ea2− (kJ mol−1) A3+ (min−1) Ea3+ (kJ mol−1) A3− (min−1) Ea3− (kJ mol−1)
Supported catalyst
Supported + dispersed catalyst
2.2E+13 194.85 0.0281 0 2.6E+6 110.45 3E+17 241.06 0 121.54
7.8E+12 194.85 0.0039 0 2.1E+5 110.44 1E+17 241.06 0 121.54
Fig. 10. Calculated effect of exothermic gain on sediment formation.
ebullated bed offsets (flow maldistribution), contact of hydrogen with liquid feed is insufficient which reduces the required amount of hydrogen to react. A similar effect is also related to lower catalyst activity. Unreacted hydrogen increases the partial pressure in the gas-liquid separator. Ebullated bed offsets and lower catalyst activity result in higher sediment formation. Exothermic temperature gain in the first reactor stage indirectly represents the intensity of hydrogenation and the content of unsaturated components in the feed. The higher the exothermic gain, the higher the extent of hydrogenation reactions until the equilibrium is reached. Fig. 10 shows the correlation between sediment formation and the exothermic temperature gain. For both regimes, the trend is decreasing; with higher hydrogenation intensity, the sediment formation is suppressed. The trend for the regime with supported catalyst only showed higher sensitivity of the sediment fraction value to the change of the exothermic temperature gain; for the regime with dispersed catalyst, the trend is almost linear and with smaller slope. The hydrogen makeup to the feed mass ratio showed positive influence on the sediment formation for both regimes as shown in Fig. 11. The amount of hydrogen make-up was maintained to compensate for the hydrogen loss in the recycle gas due to the reaction. Increased hydrogen consumption indicates higher content of aromatic and unsaturated compounds in the feed which also act as coke precursors [7].
The model shows good agreement between the calculated and the measured content of sulfur and hydrogen sulfide in the vacuum residue. Higher discrepancies were observed for the vacuum gas oil sulfur content for the regime with dispersed catalyst, which is probably caused by the higher variability of vacuum gas oil cracking and hydrodesulfurization and the relatively low sulfur content. Parity plots are shown in Figs. 12 and 13. Average mass fraction of sulfur in the unconverted vacuum
4.3. HDS performance analysis The proposed HDS model considers three forms of sulfur: sulfur content in the vacuum residue fraction, sulfur content in the vacuum gas oil fraction and hydrogen sulfide. For each regime, a separate set of kinetic parameters was estimated (Table 2). Activation energies were the same for both regimes, thus rate constant is depended on the preexponential factor. Zero preexponential factor was estimated for ks3− for both regimes, which means that the model considers hydrodesulfurization of the vacuum gas oil as an irreversible process. Zero activation energy was estimated for ks2+ for both regimes, which means that the rate of vacuum residue hydrodesulfurization does not depend on temperature.
Fig. 11. Calculated effect of hydrogen makeup to the feed mass ratio on sediment formation.
Fig. 12. Parity plot for vacuum residue and vacuum gas oil sulfur content calculation; series marked with - disp belongs to the regime with dispersed catalyst.
Fig. 13. Parity plot for hydrogen sulfide calculation; series marked with - disp belongs to the regime with dispersed catalyst.
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Fig. 14. Temperature effect of the HDS process calculated by the model at constant residence time of 185 min; series marked with - disp belongs to the regime with dispersed catalyst.
residue for the regime without dispersed catalyst was 0.00447. The addition of dispersed catalyst decreased this content to 0.00378. The HDS model was used to calculate the temperature effect (Fig. 14) and the evolution of sulfur distribution in time (Fig. 15). Temperature effect was calculated at the constant residence time of 185 min and the initial sulfur mass fraction 0,03 in VR. Sulfur content for the regime without dispersed catalyst was analyzed at the reaction temperatures of 399–412 °C and at 412–419 °C for the regime with dispersed catalyst. Data calculated outside these ranges are considered as extrapolated. The regime without dispersed catalyst showed only slightly decreasing trend of HDS conversion with temperature (Fig. 14). In Fig. 15 the steady state can be seen 65 min of residence time with 86,5% HDS conversion achieved. However, the regime with dispersed catalyst appears in dynamic state. HDS conversion linearly rises with temperature and do not reach steady state during 300 min of residence time. Higher temperatures improved the sulfur conversion of both VR and VGO. At reaction temperatures lower than 415 °C HDS conversions is lower for the regime with dispersed catalyst, but at 417 °C, HDS conversion rises up to 87.9%. These dynamic characteristics and the higher HDS conversion can be assigned to the hydrogenation effect of the dispersed catalyst and deeper cracking conversion. 5. Conclusion A lumped model based on an industrial scale unit was developed to predict the amount of sediment formed during residual hydrocracking.
It was assumed that the main sources of sediment formation are the heaviest fraction lumps in the reaction mixture: vacuum gas oil and vacuum residue. Equation for the sediment formation calculation includes four process parameters: average bed temperature, hydrogen partial pressure after the third reactor stage, exothermic temperature gain in the first stage and hydrogen makeup to feed mass ratio. Reaction temperature, hydrogen partial pressure after the third stage and hydrogen makeup to feed ratio showed increasing effect on the sediment formation while the exothermic temperature gain showed decreasing effect on the formed sediment. Two sets of model parameters were estimated for the regime with the classical supported Ni-Mo/Al2O3 catalyst and for that with a combination of supported and dispersed Mo-based catalyst. Sediment formation was examined for both regimes. Average bed temperature shows positive effect on sediment formation for the regime with supported catalyst. This effect was not observed when dispersed catalyst was incorporated, sediment formation shows no correlation with average bed temperature. This can be explained by hydrogenating contribution of the dispersed phase catalyst which suppress radical polymerization and stabilize the feedstock even at higher temperatures. The hydrogenation effect was mathematically represented by exothermic temperature gain in the first reactor stage, as in first stage hydrogenation reactions are expected to be predominant. For both regimes, sediment formation has decreasing trend with exothermic temperature gain. This supports assumption that hydrogenation suppress sediment formation. Aromatic and unsaturated compounds in feed, which act as coke precursors, are indicated by hydrogen consumption. The hydrogen makeup to the feed mass ratio showed, as expected, positive influence on the sediment formation for both regimes. Addition of the dispersed phase catalyst decreased the average as well as the median sediment fraction from 0,28 mass% to 0,24 mass% even under higher severity conditions. A hydrodesulfurization model with sets of model parameter for both regimes was also developed to predict sulfur distribution in form of hydrogen sulfide, vacuum gas oil sulfur and unconverted vacuum residue sulfur. Average HDS conversion for the regime without dispersed was 86.5% and shown slightly decreasing trend with reaction temperature. On contrary, when dispersed catalyst was employed, HDS conversion rises with reaction temperature. At reaction temperature 417 °C the HDS conversion was 87.9%. Acknowledgement This work was supported by the Grant APVV-15-0148 provided by the Slovak Research and Development Agency. The authors would like to thank Robert Žajdlík, Juraj Sláva and Michal Báhidský of Slovnaft, a. s. for providing valuable comments.
Fig. 15. Evolution of sulfur distribution in time calculated by the model at constant reaction temperature of 410 °C; series marked with - disp belongs to the regime with dispersed catalyst.
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Fuel Processing Technology journal homepage: www.elsevier.com/locate/fuproc
Research article
Hydrocracking of vacuum residue with solid and dispersed phase catalyst: Modeling of sediment formation and hydrodesulfurization Eduard Manek, Juma Haydary ⁎ Institute of Chemical and Environmental Engineering, Faculty of Chemical and Food Technology, Slovak University of Technology, Radlinského 9, 812 37 Bratislava, Slovakia
a r t i c l e
i n f o
Article history: Received 13 December 2016 Received in revised form 31 January 2017 Accepted 5 February 2017 Available online xxxx Keywords: Residual hydrocracking Sediment formation Hydrodesulfurization Mathematical modeling
a b s t r a c t Effect of dispersed Mo based catalyst on residual hydrocracking was investigated in this paper. The study was performed in an industrial scale hydrocracking unit with the average reactor temperature of 399–419 °C and pressure of 18 MPa. Vacuum residue from Ural crude oil was fed to the reactor. A sediment formation model was proposed with parameters sets for the regime with classical supported Ni-Mo/Al2O3 catalyst and for that with a combined system of supported and dispersed catalyst. Four process parameters were incorporated in the model equation: reaction temperature, hydrogen partial pressure, exothermic gain and hydrogen makeup to feed mass ratio. A decrease of the sediment formation with the application of dispersed Mo catalyst was observed, which results in higher Mo concentration in the liquid phase and thus in higher hydrogenation activity and stabilization of coke precursors. Hydrodesulfurization was also investigated and a sulfur distribution model was created with parameters for both regimes. The model contains three lumps with significant sulfur content: sulfur contained in the vacuum residue, sulfur contained in the vacuum gas oil and hydrogen sulfide. Application of the catalyst and increase of reaction temperature from 399–414 °C to 414–419 °C resulted in increase of average sulfur conversion in vacuum residue from 86.5% to 87.9%. © 2017 Published by Elsevier B.V.
1. Introduction Hydrocracking is a well-established refining process for upgrading heavy feedstock into valuable products such as gasoline, kerosene and diesel oil. First commercial hydrocracking units operation began during the interwar period [1], many enhancements have been proposed since then. One of the current challenges is to utilize the growing share of heavy, extra-heavy and alternative feedstocks due to the depletion of light petroleum reserves [2,3]. These materials contain high amounts of metals (Ni, V), organo-sulfur compounds and asphaltenes which cause operational problems. Classical solid hydrocracking catalysts cannot effectively process these impurities, and the application of dispersed phase catalysts draws more attention. Several studies and reviews dealing with slurry or dual phase catalysis (the use of both supported and dispersed catalysts) of heavy residue hydroconversion have been published in the past few years [4–9]. Active component of these dispersed phase catalysts is a metal sulfide (mostly MoS2) generated in situ by thermal decomposition of the precursor (metallic oxides, organo-metallic compounds). Panariti et al. studied the performance of different oil-soluble precursors and introduced the following order of activity Mo N Ni ~ Ru N Co N V N Fe. Advantages of dispersed catalysts over supported ones is their lower particle size, better mixing in the reaction ⁎ Corresponding author. E-mail address: [email protected] (J. Haydary).
http://dx.doi.org/10.1016/j.fuproc.2017.02.003 0378-3820/© 2017 Published by Elsevier B.V.
system, reduced hot spots and concentration gradient, promotion of hydrogen transfer, which also stabilizes the mixture and suppresses excessive cracking and coke formation [4]. In combination with an acid catalyst, the presence of dispersed MoS2 protects the main cracking catalyst long enough thus allowing the exploitation of unconventional resources [5]. The main limitation in the commercialization of this application is the high investment and operating cost caused by catalyst makeup or complicated catalyst recovery. Bacaud, in his review [9], summarized the expected properties and criteria for a convenient dispersed catalyst as follows: hydrogenation catalyst, active phase is stable under H2S/H2 pressure, high intrinsic activity, low propensity to agglomeration, sintering and sedimentation, cheap and easily disposable. Coke and sediment formation as one of the main bottlenecks of residual oil hydrocracking has drawn the attention of both academic and corporate researchers. Carbonaceous deposits can deactivate the catalyst by poisoning of the active sites or by pore blockage [10]. Retention of carbonaceous molecules on the catalysts is mainly due to their strong adsorption and low volatility (gas-phase reactions) or low solubility (liquid-phase reactions). At high temperatures (N 350 °C), the coke components are mainly polyaromatics. Their formation involves also hydrogen transfer (acid catalysts) and dehydrogenation (bifunctional catalysts) steps in addition to the condensation and rearrangement steps. On microporous catalysts, the retention of coke molecules is caused by their steric blockage within the micropores [11]. Tailleur et al. [12] studied pore plugging effects in a hydrocracking catalyst in an
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ebullated bed reactor. The obtained results indicated that the initial deactivation is mainly due to coke depositions, while its impact on the mass transfer reaction control depends on temperature. In long-term deactivation, metal deposition plays a more important role in blocking the internal micro- and meso-structures and in the formation of an external layer of pellets. Besides catalyst deactivation, coke sedimentation, which occurs in all heavy oil processing equipment, leads to other various problems [13–16]. Fouling in a heat exchanger network reduce the amount of heat recovered from product streams, which have to be compensated by higher furnace fuel consumption and subsequently higher deterioration of the furnace equipment. As the maximum heating capacity or hydraulic limit (due to pipe plugging) in the heat exchanger network is reached, throughput of the whole unit is reduced or in some cases the unit has to be shut down. In the highly competitive environment of petroleum industry, unplanned shutdowns have unfavorable consequences. High fouling rates lead to increased equipment cleaning and cost. The most common mechanism of carbonaceous deposits is free radical polymerization, where unsaturated hydrocarbons react to form longer chain molecules. This mechanism can be inhibited by hydrogenation conditions or by polymerization inhibitors stabilizing the free radicals and terminating propagation reactions [14]. Condensation reactions and asphaltene precipitation also contribute to the sediment deposition. Stanislaus et al. [15] analyzed sediments produced during VR hydrocracking in a pilot plant test. The analyzed sediments contained mainly organic materials, large part of it being sufficiently volatile to boil in the same range as heavy gas oil. A large portion (about 70 wt%) of the sediment is soluble in various solvents. Common practice is to continually apply a dispersant or solvents to prevent particle agglomeration and the formation of large clusters which are deposited more easily. Storm et al. [16] observed sediment formation during laboratory hydrotreating experiments of vacuum residue. Correlations were found between the amount of sediment produced per unit of weight of the vacuum residue feed and four chemical characteristics of the vacuum residue: degree of condensed polynuclear aromaticity, average number of alkyl-groups substituting the polynuclear aromatics, ratio of heptane insolubles to pentane insoluble-heptane solubles, and H/C ratio of the latter fraction. Majority of papers dealing with heavy oil hydrocracking provides experimental data from laboratory scale reactor or pilot plants. Modeling based on industrial scale units suffers from obvious complications and limitations. Process parameters like the reactor temperature are limited by coke generation and catalyst deactivation. Step tests represent an economical burden. Measuring instruments are prone to decalibration during a long run. Composition of the feed may vary over time due to the composition of the crude oil supplied or the performance of adjacent distillation units. Large volumes of the pipeline, columns, reactor, and separator vessels cause long delays between the changes of process parameters and the measured output. For large continuous processes it is also much more difficult to determine the exact reaction time than for small batch processes. Kinetic models were applied to optimize industrial scale hydrocracker units [17–20]. Sildir et al. [17] developed a continuous non-isothermal model for an industrial hydrocracker unit with the ability of real time optimization and online model parameter tuning. The reactor model was later extended with an additional fractionator model and a cascade model predictive control structure was developed to operate both the reactor and the fractionation column at maximum profits [18]. Zhou et al. [20] used ASPEN PLUS to predict feedstock properties and applied a lumped model to determine the optimal reactor bed temperatures for maximal diesel and kerosene yields. Such model also incorporates simplified HDN and HDS kinetics with nitrogen and sulfur as single lumps. In a previous paper [21], a lumped model of an industrial scale residual hydrocracker unit was developed. The model was implemented in the ASPEN PLUS environment to simulate the fractionation section. Simulated product distillation curves showed good
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agreement with the results of plant analysis. Main objectives of the present work are: – find the correlation between sediment formation and process parameters – extend the existing reaction network with a sediment lump and to estimate the model parameters – compare sediment formation for regimes with and without dispersed catalyst – develop a multi-lump HDS model and compare the HDS performance for regimes with and without dispersed catalyst Nomenclature [] mass fraction HDS hydrodesulfurization RHC residual hydrocracker LCO light cycle oil MCB main column bottom FCC fluid catalytic cracker sed sediments preexponential factor, min−1 Ai Eai activation energy, kJ mol−1 G offgas GLN naphtha GO gas oil Ke kerosene kinetic constant, min−1 ki R gas constant, J mol−1 K−1 t reaction (retention) time, min T temperature, °C VGO vacuum gas oil VR vacuum residue w mass fraction n9, n10, nH2, nExo, nφ model parameters P overall pressure, kPa hydrogen partial pressure, kPa PH2 exothermic temperature gain, °C Texo first stage average bed temperature, °C TR1 Ф hydrogen makeup to feed mass ratio sulfur content in vacuum residue (unc - unconverted) SVR sulfur content in vacuum gas oil SVGO 2. Experimental data All data in this work were collected from an industrial scale residual hydrocracker unit (RHC). Process parameters (flow, temperature, pressure) were measured by standard industrial measuring instruments and physicochemical (sulfur content - STN EN ISO 8754, sediments - ASTM D4870) properties were analyzed by laboratory methods. The feed for the residual hydrocracker unit is the vacuum residue from primary distillation of Ural crude oil. Vacuum residue is the heaviest fraction obtained from primary crude oil distillation, which contains substances with high molecular weight (about 600 × kg kmol−1), its distillation curve is in the range of 440 + °C, and the density is about 1020 kg m− 3 at 20 °C. High sulfur content (about 3.00 mass%), nitrogen (0.45 mass%), and metals (vanadium: 210 mg kg−1, nickel: 66.5 mg kg−1, and sodium: 4.35 mg kg−1). The feed is mixed with heavy aromatic oils, light cycle oil (LCO) and the main column bottom (MCB) from the FCC unit to suppress coagulation of asphaltenes, sediment formation and to maintain required hydrodynamic characteristics of the feed. Diluents represent almost 12 mass% of the whole feed mixture. Before entering the reactor cascade, preheated feed is mixed with high temperature hydrogen. The industrial reactor cascade (Fig. 1) consisted of three identical ebullated bed reactors (diameter of 3.6 m, height of 36.5 m, ebullated bed of approximately 28 m) with a commercial Ni-Mo/Al2O3 hydrocracking catalyst (pellets with a 1 mm diameter and 3.2 mm length).
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Fig. 1. Cascade of ebullated bed reactors.
Temperature in the reactors was held at around 401–419 °C and the pressure at 18–20 MPa. Effluents from the first to the second stage and from the second to the third stage were quenched by hydrocarbon oil and hydrogen to maintain the temperature of the ebullated beds constant and also to compensate for spent hydrogen. Hydrogen makeup is controlled based on the pressure set point in the 1st stage suction drum of a four stage compressor. In the 2nd, 3rd and 4th stage, both hydrogen make-up and recycle gas from the gas-liquid separator are compressed. The process includes on-stream catalyst addition and withdrawal avoiding the reactor shut down for catalyst replacement, which is critical because the catalyst bed is continuously deactivated by metal adsorption and coke deposits. The ebullated bed reactor system is comprehensively described by Martínez et al. [22]. The product stream from the third reactor enters a high pressure separator, where liquid and gas phase are separated. Both streams are subsequently cooled, depressurized and separated into final product fractions in an atmospheric and a vacuum column. Except for the classical solid Ni-Mo/Al2O3 catalyst, an additional commercial Mo-based dispersed phase catalyst was tested with solid catalyst present in the ebullated bed. Oil soluble precursor were injected to the LCO stream at a constant rate, overall mass content of catalyst in the feed was 0.021 mass%. The main purpose of this catalyst is to promote hydrogenation, asphaltene stabilization, to upgrade of the product quality, lower the downstream equipment fouling and to enable operation at higher reactor temperatures without excessive coke formation. Sediments were analyzed in the unconverted residue via an ASTMD4870. Analyzed sediments consisted of both organic and inorganic types. The inorganic part consisted of rust, sand, and crushed catalyst particles; the organic sediments are formed of coke and asphaltenes. Hydrotreating affected only the formation of the organic part; it can be assumed that only a negligible amount of inorganic sediments formed. Sulfur content in both feed and products was analyzed. A significant amount of sulfur was present in the vacuum residue feed, unconverted vacuum residue and in the vacuum gas oil. Sulfur content in other products was in ppm and it was not included in the mathematical model.
vacuum residue and vacuum gas oil mass fraction. d½VR ¼ −ðk1 þ k2 þ k3 þ k4 þ k7 Þ½VR dt
ð1Þ
d½VGO ¼ k4 ½VR−ðk5 þ k6 þ k8 Þ½VGO dt
ð2Þ
d½GO ¼ k3 ½VR þ k5 ½VGO dt
ð3Þ
d½Ke ¼ k7 ½VR þ k8 ½VGO dt
ð4Þ
d½GLN ¼ k2 ½VR þ k6 ½VGO dt
ð5Þ
d½G ¼ k1 ½VR dt
ð6Þ
3.1. Sediment formation model
P H2 nH2 T Exo nExo nφ d½sed : :Ф ¼ k9 :½VRn9 þ k10 :½VGOn10 : P T R1 dt
ð7Þ
Additional process parameters were included in Eq. (7) for the determination of the sediment formation rate which is based on the mass fraction of the vacuum residue ([VR]) and the vacuum gas oil ([VGO]). Initial value of the sediment mass fraction ([sed]) of 0 and irreversible reaction were assumed. Hydrogen partial pressure (PH2) was
3. Mathematical modeling The lumped model of vacuum residue hydrocracking proposed in our previous work consists of six fraction lumps: vacuum residue - VR, vacuum gas oil - VGO, gas oil - GO, kerosene - Ke, gasoline - GLN, gas – G, and of eight reaction steps (Eqs. (1)–(6)). The model was upgraded to predict sediment formation. It was assumed that sources of organic sediments are the vacuum residue and vacuum gas oil lumps. The upgraded reaction network is shown in Fig. 2. Pathways k9 and k10 were not included in Eqs. (1) and (2) due to insignificant effect of
Fig. 2. Reaction network of vacuum residue hydrocracking [21] with two additional steps of sediment formation.
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measured based on the hydrogen concentration and the overall pressure (P) in the gas/liquid separator after the third reactor stage. Partial pressure represents the remaining hydrogen concentration in the gas phase at the end of the reactor cascade. Higher amount of hydrogen should suppress the sediment formation and asphaltene stabilization. Exothermic temperature gain, TExo, across the ebullated bed in the first reactor stage indicated the intensity of hydrogenation; this value is the highest due to the initial hydrogenation reactions in first stage and therefore it was chosen as an equation parameter. To provide dimensionless values, TExo figures in Eq. (7) as ration with average temperature in first stage ebullated bed. Symbol Ф represents the hydrogen makeup to the feed mass ratio before the first reactor stage. Kinetic constants k9 and k10 for sediment formation from the vacuum residue and the vacuum gas oil, respectively, were calculated using the Arrhenius Eq. (8): Eai
ki ¼ Ai :e− RT
ð8Þ
Index i represents the respective reaction pathway. Nine model parameters: A9, A10, Ea9, Ea10, n9, n10, nH2, nExo, nφ, were estimated. Two sets of parameters for both regimes with and without an addition of dispersed phase catalyst were estimated. Objective function was the sum of least squares of real and calculated sediment values. 3.2. Sulfur distribution model In our previous work [21], a simplified sulfur conversion model was proposed. Although this model provides a satisfactory agreement between the calculated and real sulfur conversion, the new model was designed to calculate sulfur distribution considering hydrogen sulfide and the remaining amount in the vacuum residue and the vacuum gas oil fraction, other products contain only an insignificant amount of sulfur. A scheme of the reaction path is shown in Fig. 3. SVR, SVGO and SH2S represents sulfur mass content in vacuum residue, mass content in vacuum gas oil and sulfur mass in hydrogen sulfide. Step ks1 represents the transfer of vacuum residue sulfur to the vacuum gas oil sulfur by an irreversible cracking mechanism. Reversible hydrodesulfurization of the vacuum residue and vacuum gas oil is represented by steps ks1, ks2 −, ks2 +, ks3 −, ks3+. Activation energies and preexponential factors of the Arrhenius Eq. (8) were estimated by minimizing the sum of least squares of real and calculated sulfur content in the vacuum residue and the vacuum gas oil. The amount of sulfur in form of hydrogen sulfide was calculated using the mass balance [Eq. (9)]. Residence time was approximately calculated from the volumetric flow of the feed and the ebullated bed volume. unc Sfeed VR ¼ SVR þ SVGO þ SH2S
Fig. 4. Formation of sediments without dispersed Mo catalyst.
4. Results and discussion 4.1. Analysis of sediment formation Formation of sediments measured during the vacuum residue hydrocracking without the application of dispersed catalyst for over 130 day period is shown in Fig. 4. A 200 day process employing the dispersed catalyst is shown in Fig. 5. A moving average was applied to both calculated and experimental sediment values to eliminate scattering and highlight their trend. The average sediment content in the unconverted vacuum residue without dispersed phase catalyst was 0.28 mass% and the median value was 0.26 mass%, operated in lower severity regime with the average bed temperature of 400–413 °C. The addition of the dispersed phase catalyst decreased the average and the median sediment fraction to 0.24 mass% even in higher severity regime with the average bed temperature of 413–419 °C. A few discrepancies between the real and the predicted sediment values were observed. Inconsistent peaks of experimental and calculated values were visible at days 17, 62, 93, 132, 151 and 180 (Fig. 4) and at days 16 and 127 (Fig. 5), probably due to delayed sediment accumulation and purging from the equipment and pipeline surface, changes of the feedstock quality and other unregistered process deviations. The sediment drop at day 75 for the supported catalyst only regime (Fig. 4), correlates with the rapid decrease of the reactor temperature from 410 °C to 400 °C. Re-establishing the temperature of 412 °C led to an immediate increase of the sediment content. The sediment peak at day 105 for the regime with dispersed catalyst correlates with the increase of the hydrogen partial
ð9Þ
Fig. 3. Reaction scheme for hydrodesulfurization: vacuum residue sulfur content (SVR) and vacuum gas oil sulfur content (SVGO).
Fig. 5. Formation of sediments with dispersed Mo catalyst.
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E. Manek, J. Haydary / Fuel Processing Technology 159 (2017) 320–327 Table 1 Estimated parameters for Eq. (7).
A9 (min−1) Ea9 (kJ mol−1) A10 (min−1) Ea10 (kJ mol−1) n9 n10 nH2 nExo nφ
Supported catalyst
Supported + dispersed catalyst
3.92E+17 242.8 1.05E+23 141.8 476.4 70.52 2.4507 −1.758 6.817
5135 0 1.05E+23 141.8 152.4 175.1 1.456 −1.543 0.9998
4.2. Effect of process parameters on sediment formation
Fig. 6. Parity plot for measured and calculated sediment data.
pressure peak. It is obvious, that the model parameters were estimated based on the degree of change of individual process parameters and the subsequent change of the sediment fraction. Model parameters can be tuned after every non-standard process parameter change to keep the model updated and include new observed effects. A comparison of measured and calculated sediments by a parity plot is given in Fig. 6. Residual plot of measured experimental data is presented in Fig. 7. Regression coefficient for supported catalyst only and combined regime (supported + dispersed) is 0.96 and 0.98, respectively. Residuals are distributed symmetrically in both regimes without any clear pattern. Estimated model parameters for both regimes are shown in Table 1. Due to the very low content of sediments (0.1–0.5 mass%) and the parametric form of Eq. (7), kinetic parameters cannot be compared with values presented in literature. Parameter nH2 for partial hydrogen pressure is almost zero for the regime without dispersed catalyst, which means that the partial hydrogen pressure does not affect the sediment calculation for this regime. Model parameters in this work are plant specific (specific equipment setup, catalyst and feed properties); however, the trend introduced in this study should describe residual hydrocracking in general.
The effects of average bed temperature, hydrogen partial pressure, exothermic temperature gain in the first reactor stage and the hydrogen makeup to feed mass ratio on the sediment were calculated using the presented model. One process parameter was varied while other parameters were set constant at median values. The sediment fraction rapidly increased with the increasing reaction temperature (Fig. 8). Higher temperature intensifies the cracking reactions as well as the radical mechanism of coke formation. Model of the regime employing both supported and dispersed phase catalysts did not show this trend probably due to the promotion of hydrogenation reactions by the dispersed phase catalyst which suppress radical polymerization and stabilize the feedstock. Partial hydrogen pressure in the separator after the third reactor stage showed a significant effect on the sediment formation (Fig. 9) for the regime with both supported and dispersed catalysts. For ideal operation, partial pressure is related to hydrogen make-up. During
Fig. 8. Calculated effect of average bed temperature on sediment formation.
Fig. 7. Residual plot for measured sediment data.
Fig. 9. Calculated effect of hydrogen partial pressure after third stage on sediment formation.
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Table 2 Estimated kinetic parameters for the HDS model.
A1 (min−1) Ea1 (kJ mol−1) A2+ (min−1) Ea2+ (kJ mol−1) A2− (min−1) Ea2− (kJ mol−1) A3+ (min−1) Ea3+ (kJ mol−1) A3− (min−1) Ea3− (kJ mol−1)
Supported catalyst
Supported + dispersed catalyst
2.2E+13 194.85 0.0281 0 2.6E+6 110.45 3E+17 241.06 0 121.54
7.8E+12 194.85 0.0039 0 2.1E+5 110.44 1E+17 241.06 0 121.54
Fig. 10. Calculated effect of exothermic gain on sediment formation.
ebullated bed offsets (flow maldistribution), contact of hydrogen with liquid feed is insufficient which reduces the required amount of hydrogen to react. A similar effect is also related to lower catalyst activity. Unreacted hydrogen increases the partial pressure in the gas-liquid separator. Ebullated bed offsets and lower catalyst activity result in higher sediment formation. Exothermic temperature gain in the first reactor stage indirectly represents the intensity of hydrogenation and the content of unsaturated components in the feed. The higher the exothermic gain, the higher the extent of hydrogenation reactions until the equilibrium is reached. Fig. 10 shows the correlation between sediment formation and the exothermic temperature gain. For both regimes, the trend is decreasing; with higher hydrogenation intensity, the sediment formation is suppressed. The trend for the regime with supported catalyst only showed higher sensitivity of the sediment fraction value to the change of the exothermic temperature gain; for the regime with dispersed catalyst, the trend is almost linear and with smaller slope. The hydrogen makeup to the feed mass ratio showed positive influence on the sediment formation for both regimes as shown in Fig. 11. The amount of hydrogen make-up was maintained to compensate for the hydrogen loss in the recycle gas due to the reaction. Increased hydrogen consumption indicates higher content of aromatic and unsaturated compounds in the feed which also act as coke precursors [7].
The model shows good agreement between the calculated and the measured content of sulfur and hydrogen sulfide in the vacuum residue. Higher discrepancies were observed for the vacuum gas oil sulfur content for the regime with dispersed catalyst, which is probably caused by the higher variability of vacuum gas oil cracking and hydrodesulfurization and the relatively low sulfur content. Parity plots are shown in Figs. 12 and 13. Average mass fraction of sulfur in the unconverted vacuum
4.3. HDS performance analysis The proposed HDS model considers three forms of sulfur: sulfur content in the vacuum residue fraction, sulfur content in the vacuum gas oil fraction and hydrogen sulfide. For each regime, a separate set of kinetic parameters was estimated (Table 2). Activation energies were the same for both regimes, thus rate constant is depended on the preexponential factor. Zero preexponential factor was estimated for ks3− for both regimes, which means that the model considers hydrodesulfurization of the vacuum gas oil as an irreversible process. Zero activation energy was estimated for ks2+ for both regimes, which means that the rate of vacuum residue hydrodesulfurization does not depend on temperature.
Fig. 11. Calculated effect of hydrogen makeup to the feed mass ratio on sediment formation.
Fig. 12. Parity plot for vacuum residue and vacuum gas oil sulfur content calculation; series marked with - disp belongs to the regime with dispersed catalyst.
Fig. 13. Parity plot for hydrogen sulfide calculation; series marked with - disp belongs to the regime with dispersed catalyst.
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Fig. 14. Temperature effect of the HDS process calculated by the model at constant residence time of 185 min; series marked with - disp belongs to the regime with dispersed catalyst.
residue for the regime without dispersed catalyst was 0.00447. The addition of dispersed catalyst decreased this content to 0.00378. The HDS model was used to calculate the temperature effect (Fig. 14) and the evolution of sulfur distribution in time (Fig. 15). Temperature effect was calculated at the constant residence time of 185 min and the initial sulfur mass fraction 0,03 in VR. Sulfur content for the regime without dispersed catalyst was analyzed at the reaction temperatures of 399–412 °C and at 412–419 °C for the regime with dispersed catalyst. Data calculated outside these ranges are considered as extrapolated. The regime without dispersed catalyst showed only slightly decreasing trend of HDS conversion with temperature (Fig. 14). In Fig. 15 the steady state can be seen 65 min of residence time with 86,5% HDS conversion achieved. However, the regime with dispersed catalyst appears in dynamic state. HDS conversion linearly rises with temperature and do not reach steady state during 300 min of residence time. Higher temperatures improved the sulfur conversion of both VR and VGO. At reaction temperatures lower than 415 °C HDS conversions is lower for the regime with dispersed catalyst, but at 417 °C, HDS conversion rises up to 87.9%. These dynamic characteristics and the higher HDS conversion can be assigned to the hydrogenation effect of the dispersed catalyst and deeper cracking conversion. 5. Conclusion A lumped model based on an industrial scale unit was developed to predict the amount of sediment formed during residual hydrocracking.
It was assumed that the main sources of sediment formation are the heaviest fraction lumps in the reaction mixture: vacuum gas oil and vacuum residue. Equation for the sediment formation calculation includes four process parameters: average bed temperature, hydrogen partial pressure after the third reactor stage, exothermic temperature gain in the first stage and hydrogen makeup to feed mass ratio. Reaction temperature, hydrogen partial pressure after the third stage and hydrogen makeup to feed ratio showed increasing effect on the sediment formation while the exothermic temperature gain showed decreasing effect on the formed sediment. Two sets of model parameters were estimated for the regime with the classical supported Ni-Mo/Al2O3 catalyst and for that with a combination of supported and dispersed Mo-based catalyst. Sediment formation was examined for both regimes. Average bed temperature shows positive effect on sediment formation for the regime with supported catalyst. This effect was not observed when dispersed catalyst was incorporated, sediment formation shows no correlation with average bed temperature. This can be explained by hydrogenating contribution of the dispersed phase catalyst which suppress radical polymerization and stabilize the feedstock even at higher temperatures. The hydrogenation effect was mathematically represented by exothermic temperature gain in the first reactor stage, as in first stage hydrogenation reactions are expected to be predominant. For both regimes, sediment formation has decreasing trend with exothermic temperature gain. This supports assumption that hydrogenation suppress sediment formation. Aromatic and unsaturated compounds in feed, which act as coke precursors, are indicated by hydrogen consumption. The hydrogen makeup to the feed mass ratio showed, as expected, positive influence on the sediment formation for both regimes. Addition of the dispersed phase catalyst decreased the average as well as the median sediment fraction from 0,28 mass% to 0,24 mass% even under higher severity conditions. A hydrodesulfurization model with sets of model parameter for both regimes was also developed to predict sulfur distribution in form of hydrogen sulfide, vacuum gas oil sulfur and unconverted vacuum residue sulfur. Average HDS conversion for the regime without dispersed was 86.5% and shown slightly decreasing trend with reaction temperature. On contrary, when dispersed catalyst was employed, HDS conversion rises with reaction temperature. At reaction temperature 417 °C the HDS conversion was 87.9%. Acknowledgement This work was supported by the Grant APVV-15-0148 provided by the Slovak Research and Development Agency. The authors would like to thank Robert Žajdlík, Juraj Sláva and Michal Báhidský of Slovnaft, a. s. for providing valuable comments.
Fig. 15. Evolution of sulfur distribution in time calculated by the model at constant reaction temperature of 410 °C; series marked with - disp belongs to the regime with dispersed catalyst.
E. Manek, J. Haydary / Fuel Processing Technology 159 (2017) 320–327
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