Comparison of the Performance of Upflow and Downflow Small

Energy & Fuels 2001, 15, 1139-1144

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Comparison of the Performance of Upflow and Downflow Small-Reactors in Hydrodesulfurization Reactions L. C. Castan˜eda-Lo´pez,† F. Alonso-Martı´nez,† J. Ancheyta-Jua´rez,*,†,‡ S. K. Maity,† E. Rivera-Segundo,† and M. N. Matus-Guerra† Instituto Mexicano del Petro´ leo, Eje Central La´ zaro Ca´ rdenas 152, 07730 Me´ xico, D.F., Mexico, and Instituto Polite´ cnico Nacional, ESIQIE, 07738 Me´ xico, D.F., Mexico Received January 8, 2001. Revised Manuscript Received June 12, 2001

An experimental study for comparing the behavior of a small-reactor in the hydrodesulfurization of straight-run gas oil with two modes of operation, upflow and downflow, is presented. The reactor has 525 mm length and 14.3 mm internal diameter. Experiments were conducted at commercial HDS operating conditions (5.3 MPa total pressure, 356.2 mL/mL hydrogen-to-oil ratio, reaction temperature of 613-653 K, and LHSV of 1.0-2.5 h-1). All tests were carried out over a commercial NiMo/γ-Al2O3 catalyst. The effect of different catalyst-to-diluent ratios on the sulfur removal was also studied. The results of SRGO hydrodesulfurization showed that upflow with a 50/50 vol % catalyst-to-diluent ratio is the best system to avoid external gradients and other effects such as poor wetting of catalyst, axial dispersion, wall effects, and flow maldistribution. Apparent activation energies were determined for upflow and downflow modes of operation with and without diluent.

Introduction Hydrodesulfurization (HDS) is extensively used in petroleum refining for the removal of sulfur and other heteroatoms. The HDS process is frequently employed for treating different crude oil fractions, such as straightrun naphtha and gas oil, FCC feedstock, and residuals. Depending on the feedstock and reactor conditions, the reaction can be carried out in two or three phases. In the case of SRGO, hydrodesulfurization is commonly accomplished in trickle-bed reactors (gas-liquid, downflow, fixed-bed systems). Due to the inherent economic benefits, hydrodesulfurization reactions are usually conducted in small-scale reactors for different research purposes such as catalyst screening, evaluation of alternative feedstocks, kinetic parameters estimation, studies about the effect of operating conditions on product selectivities, prediction of the behavior of industrial-scale reactors, etc.1 These laboratory reactors can be operated with upflow or downflow. Downflow or trickle-bed reactors (TBR) consist of a fixed bed of catalyst particles contacted by a cocurrent downward gas-liquid flow carrying both reactants and products. When the gas and liquid are fed cocurrently upward through the catalyst bed, the system is called a flooded-bed reactor (FBR) or upflow reactor.2 * Author to whom correspondence should be addressed. Fax: (+52-5) 333 8429. E-mail: [email protected]. † Instituto Mexicano del Petro ´ leo. ‡ Instituto Polite ´ cnico Nacional. (1) Bej, S. K.; Dabral, R. P.; Gupta, P. C.; Mittal, K. K.; Sen, G. S.; Kapoor, V. K.; Dalai, A. K. Energy Fuels 2000, 14, 701. (2) Al-Dahhan, M. H.; Larachi, F.; Dudukovic, M. P.; Laurent, A. Ind. Eng. Chem. Res. 1997, 36, 3292.

In both TBR and FBR systems, commercial catalyst samples and real feedstocks are frequently used for conducting different experimental studies; however, the length of the catalyst bed and hence the reactor length to catalyst particle diameter ratio are low as compared to commercial reactors. In addition, low liquid velocities are used in smallscale reactors in order to match the liquid hourly space velocities (LHSV) of industrial units. These differences cause a number of problems in testing catalyst having commercially applied size and shape, such as poor wetting of catalyst, wall effect, axial dispersion, maldistribution, and the data obtained in such systems may not be reliable, because they may not be totally the result of catalyst activity.1,3 The use of small catalyst particles is helpful for reducing these phenomena, this means that the commercial catalyst sample has to be crushed in order to test it in the form of fine particles. However, the data generated with crushed catalyst may not be representative for industrial practice, since pore diffusion limitation, which may occur with commercial catalyst size, will exhibit a higher activity with crushed sample.4 The effective solution that is recommended and accepted to overcome the problems in testing a commercial catalyst in small reactors is to use the catalyst in its original form but diluted with nonporous inert particles. Various researchers have studied the use of diluents in small reactors and they have attemped to overcome these problems.5-10 However, most of these studies have (3) Wu, Y.; Khadilkar, M. R.; Al-Dahhan, M. H.; Dudukovic, M. P. Ind. Eng. Chem. Res. 1996, 35, 397. (4) Sie, S. T. AIChE J. 1996, 42 (12), 3498.

10.1021/ef010006p CCC: $20.00 © 2001 American Chemical Society Published on Web 08/04/2001

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Figure 1. Diagram of the HDS test unit.

been restricted to bench scale reactors where about 50100 mL of catalysts were used. The problem with these systems is that new catalyst formulations are not available in large quantity, and the amount of catalyst involved in experimental tests must be small.11 On the other hand, only a few studies are available in the literature providing information on the effect of diluent on a small-reactor performance.1,12 As it was mentioned before, upflow and downflow reactors are often used in laboratory-scale studies; however, few comparative studies between these two systems are reported in the literature. Dudukovic et al.3 summarized some of these reports. They have found that the upflow reactor behavior can be better than downflow reactor and vice versa depending on gas and liquid velocities, level of conversion, and the type of reaction system. They have also concluded that there is no clear guidance as to which reactor will perform better for a given reaction system. To understand the differences between upflow and downflow reactors for a given reaction, more systematic studies need to be done. For this reason, in this work we present a comparative study in the hydrodesulfurization of straight-run gas oil using a small-reactor with downflow and upflow modes of operation, with different catalyst-to-diluent ratios. Experimental Section Materials. The feedstocks used in this study for hydrodesulfurization reactions were straight-run gas oils recovered from (5) Carruters, J. D.; DiCamillo, J. D. Appl. Catal. A 1998, 43, 253. (6) Al-Dahhan, M. H.; Wu, Y.; Dudukovic, M. P. Ind. Eng. Chem. Res. 1995, 34, 741. (7) Garica, W.; Pazos, J. M. Chem. Eng. Sci. 1982, 37, 1589. (8) Diaz, R. A.; Mann, R. S.; Sambi, I. S. Ind. Eng. Chem. Res. 1993, 32, 1355. (9) Chen, Y. W.; Hsu, W. C.; Lin, C. S.; Kang, B. C.; Wu, S. T.; Leu, L. J.; Wu, J. C. Ind. Eng. Chem. Res. 1990, 29, 1831. (10) Al-Dahhan, M. H.; Dudukovic, M. P. AIChE J. 1996, 42, 2594. (11) Letourneur, D.; Bacaud, R.; Vrinat, M. Ind. Eng. Chem. Res. 1998, 37, 2662. (12) Sie, S. T. Rev. Inst. Fr. Petrol. 1991, 46, 501.

Table 1. Characterization of the Feedstock (SRGOs) feedstock specific gravity 20/4 °C elemental analysis, wt % C H O N S aromatics, wt %

SRGO-1 0.8687

SRGO-2 0.8733

85.08 13.00 0.250 0.054 1.616 39.71

85.47 12.50 0.260 0.059 1.704 44.50

Table 2. Physical and Chemical Properties of the Commercial Catalyst diameter, mm length, mm surface area, m2/g pore volume, cm3/g bulk density, g/cm3 MoO3, wt % NiO, wt % P, wt %

2.3 5.2 204 0.50 0.78 9.5 2.4 1.58

an HDS commercial unit. SRGO-1 was employed for testing hydrotreating reactions without diluent and SRGO-2 for tests with diluent. The main properties of these samples are presented in Table 1. The catalyst employed was a trilobe extrudate commercially available sample. Physical and chemical properties of this catalyst are shown in Table 2. The inert used for diluting the catalyst bed was a nonimpregnated R-alumina with particle diameter between 1.4 and 2.3 mm. Hydrogen was of 99.999% purity. Reaction System. All hydrodesulfurization experiments were performed in a small-scale test plant. A detailed scheme of this unit is shown in Figure 1. The main components of this plant are the liquid and gas feeding units, the reactor, and the product recovery section. The reactor is the central part of this plant. It is made of a stainless steel tube. The total length and internal and external diameters of the reactor are 525, 14.3, and 25.4 mm, respectively. This reactor is designed to work with both modes of operation, downflow and upflow (Figure 2). The length of the reactor is subdivided into three sections. The first section was packed with inert particles and was used

Upflow and Downflow Small-Reactors in Hydrodesulfurization

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Figure 2. HDS reactor. to heat the mixture to the desired reaction temperature and to provide a uniform feedstock distribution. The following section contained the HDS catalyst and the diluent. The exit section was also packed with inert particles. Between each section, a glass-wool filter was employed in order to improve feed distribution. These catalyst and inert distributions are also shown in Figure 2. The lengths of each section are in the following ranges: section 1, 135-185 mm; section 2, 100-205 mm; and section 3, 185-240 mm. These lengths depend on the amount of catalyst and diluent loaded in the reactor. The reactor temperature was maintained at the desired level by using a three-zone electric furnace, which provided an isothermal temperature along the active reactor section. The catalytic bed temperature was measured during the experiments by a movable axial thermocouple located inside the reactor. The range of operating conditions of the small-scale plant are: 700 MPa maximum pressure, 723 K maximum temperature, and 0.5-8.0 h-1 LHSV. Catalyst Sulfiding and HDS Experiments. After loading the catalyst and diluent, the catalytic bed was in-situ sulfided with a desulfurized naphtha contaminated with 0.6 wt % CS2 at 5.3 MPa, a hydrogen-to-oil ratio of 356.2 mL/mL, at 503 K, liquid hourly space velocity of 3 h-1, 12 h, to ensure complete catalyst presulfiding. When the presulfiding was completed, the feedstock was introduced and sulfiding feed was stopped without stopping the hydrogen flow, and the reactor temperature and other conditions were adjusted to the desired start-of-run conditions. The hydrodesulfurization of SRGO was carried out at constants reaction pressure and hydrogen-to-oil ratio without hydrogen recycle (5.3 MPa and 356.2 mL/mL, respectively). Reaction temperature was studied in the range of 613-653 K, while liquid hourly space velocities were varied in the range 1.0-2.5 h-1. Experiments were performed in the small-reactor using both modes of operation, upflow and downflow. The effect of dilution of catalyst with R-alumina was studied at the following catalyst-to-diluent ratios: 25/75 vol %, and 50/50 vol %.

Table 3. D-86 ASTM Distillation Data (in K) for SRGO-2 and Products IBP 5 vol % 10 vol % 20 vol % 30 vol % 40 vol % 50 vol % 60 vol % 70 vol % 80 vol % 90 vol % 95 vol % EBP

SRGO-2

product 1

product 2

426 555 570 583 590 595 600 605 610 616 626 634 638

432 555 569 582 589 595 600 604 610 616 625 633 638

427 555 570 582 589 595 600 605 610 617 626 633 640

Results and Discussion Tests with 100% r-Alumina. Before conducting HDS experiments, it was checked that the material used as diluent (R-alumina) was actually inert, specially for cracking reactions. Two experiments were performed at 5.3 MPa, 356.2 mL/mL H2/oil ratio, 633 K reaction temperature, and 1.0 h-1 LHSV. A distillation curve given by D-86 ASTM method was selected as an analysis test to check the possible alumina activity. If this material has cracking activity, a change in the SRGO distillation curve can be observed. Table 3 shows a comparison of the feedstock and two products ASTM distillations. As can be seen there were no significant changes in this property. IBP and EBP show little differences, but it can be considered within experimental limitations. In addition, these points are not considered as representatives in this type of experimental distillation. Moreover, gas-phase samples of the reactor’s exit were analized by a gas chromatograph. We did not find any

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Figure 3. Temperature profiles in the reactor.

compositional changes between inlet and outlet gases. These results agree with others reported in the literature.11 On the basis of these results we can affirm that R-alumina is totally inert and it can be used as a diluent for the hydrodesulfurization catalyst without any problem. Isothermicity of the Reactor. Isothermal condition along the reactor length was examined in each experiment by a movable axial thermocouple located inside the reactor (Figure 2). Figure 3 shows three typical temperature profiles obtained at 613, 633, and 653 K of reaction temperature. These profiles were obtained when only catalyst was used. The highest deviation from the desired temperature value found in these tests was about 3-4 K. When the catalytic bed was diluted with R-alumina the deviation was reduced to 1-2 K. This reduction in temperature deviation from the desired value is because temperature homogeneity in the bed is improved by using catalyst bed dilution, and the chances of temperature runaway are diminished in the case of strongly exothermic reactions, such as those carried out in hydrotreating process (heats of reaction: hydrodesulfurization, -251000 kJ/kmol, hydrodenitrogenation, -64850 kJ/kmol, hydrocracking, -41000 kJ/kmol, hydrogenation, -125520 kJ/kmol).13 It can be concluded that there are no significant temperature differences between desired and measured temperatures, specially when catalyst bed dilution is employed. Comparison of Upflow and Downflow Systems without Diluent. To compare the behavior of both reactions systems, upflow and downflow, and verify the presence of external gradients, experiments were conducted with different amounts of catalyst and reaction temperatures using SRGO-1. These experiments are based on the principle that in the absence of external gradients, the conversion must be independent of the linear velocity through the bed.14 Two series of tests were run, the first using the minimum amount of catalyst in the reactor (14 mL), and the second, with the maximum amount of catalyst (29 mL). These minimum and maximum amounts of catalyst were taken because these amounts can be kept within a suitable thermal control zone in the reactor. (13) Tarhan, M. O. Catalytic Reactor Design; McGraw-Hill: New York, 1983. (14) Perego, C.; Peratello, S. Catal. Today 1999, 52, 133.

Figure 4. Effect of upflow (dotted lines) and downflow (solid lines) systems on HDS (circles: minimum amount of catalyst; squares: maximum amount of catalyst).

In each series, LHSV is varied and two curves are traced for HDS conversion vs 1/LHSV. If the two curves overlap there are no external gradients limitations.14 The tests were done for the two modes of operation, upflow and downflow, using the same amounts of catalyst in both cases. The results of these experiments are presented in Figure 4. In all these experiments, the HDS catalyst was loaded to the reactor without R-alumina. This is reported in the literature to be the best way to compare these two systems, when the beds for upflow and downflow are identically packed, and the bed voidage is constant.3 We have plotted conversion of HDS against 1/LHSV and it is presented in Figure 4. We observe from this figure that the gap between two lines (solid) is wide in the case of the downflow system. But this gap (dotted lines) is narrow in the upflow system. These phenomena are observed in both low and high temperatures. It indicates that the catalyst is better wetted in the upflow system.12 It can also be seen from Figure 4 that the gap between two lines is being reduced with decreasing of LHSV (i.e., with increase of 1/LHSV) and it is found in both upflow and downflow systems and in both low and high temperatures. It indicates that catalyst is also better at lower LHSV. From these experimental results we can conclude that external diffusional effects in upflow system are very low, and they are higher in the downflow mode of operation. Figure 5 shows some experimental results about the effect of both reaction temperature and LHSV on SRGO hydrodesulfurization using upflow and downflow systems. The differences in sulfur removal between these two systems can be clearly seen, especially at low temperature and high space velocity. In general, HDS

Upflow and Downflow Small-Reactors in Hydrodesulfurization

Figure 5. HDS at different LHSV and reaction temperature for upflow (dotted lines) and downflow (solid lines) systems. Symbols: circles (LHSV ) 1.0) and triangles (LHSV ) 2.5).

values are higher in upflow compared to downflow system. From Figure 5 it is also observed that the two lines (solid and dotted) are merged together at 1.0 h-1 LHSV and 653 K temperature. It reveals that all types of maldistribution can be minimized at lower LHSV and higher temperature irrespective of upflow and downflow modes. This observation is very important because experimental studies should be carried out preferably in upflow system in order to avoid external gradients. Of course, additional experiments have to be performed to study internal concentration gradients, which consist of determining the conversion for particles of different size at constant space velocity. The system is under chemical kinetic control when conversions are the same.14 The better behavior found in the upflow system compared to downflow is because the latter mode of operation might result in regions of poorly wetted catalyst, and the upflow system generates greater liquid holdup.5 Effect of Dilution of Catalytic Bed on HDS Conversion. Once the upflow mode of operation was defined to be better than that of the downflow system for avoiding the problems already mentioned in testing catalyst having commercially applied size and shape, we conducted some experiments in order to study the effect of dilution of the catalytic bed on SRGO hydrodesulfurization level. Unfortunately, the amount of SRGO-1 was not enough to perform these tests, and a new feed had to be employed (SRGO-2). SRGO-2 is heavier than SRGO-1, and it has higher heteroatoms contents as can be observed in Table 1. For this reason, results obtained in experiments without catalyst bed dilution are not totally comparable with those with bed dilution. Two experiments were carried out with catalytic beds corresponding to 50/50 vol % and 25/75 vol % of catalystto-diluent ratio. Both experiments were performed with the upflow system at 5.3 MPa pressure and 356.2 mL/ mL hydrogen-to-oil ratio. Reaction temperature was studied at 613, 633, and 653 K, and LHSV at 1.0 and 2.5 h-1. The results of these tests are shown in Figure 6. In both cases the maximum amount of catalyst (29 mL) was taken in order to avoid the axial dispersion which is found in short bed length.

Energy & Fuels, Vol. 15, No. 5, 2001 1143

Figure 6. HDS at different LHSV and reaction temperature for 50/50 vol % (solid lines) and 25/75 vol % (dotted lines) catalyst-to-diluent ratios. Symbols: circles (LHSV ) 1.0), and triangles (LHSV ) 2.5).

It is well-known that incomplete catalyst wetting is substantially avoided by diluting the laboratory beds with an inert. This inert material should be nonporous and preferably smaller than the catalyst. When catalyst dilution is used, the hydrodynamics is largely dictated by the packing of inert, whereas the catalytic phenomena are governed by the catalyst particle of the same shape and size as used in industrial plants.2 The diluent used in the present study was R-alumina (1.4-2.3 mm diameter), which exhibited an average smaller particle diameter compared to the commercial NiMo catalyst. It can be clearly observed that 50/50 vol % of catalystto-diluent ratio presents higher HDS conversions than the 25/75 vol % ratio. This behavior is because the 50/ 50 vol % ratio has a beneficial effect on external concentration gradients. On the other hand, the amount of catalyst is not enough to reach good contact between feedstock and catalyst when the 25/75 vol % ratio is used. That is the main reason for having low HDS conversions with this latter catalyst-to-diluent ratio. Finally, we can state that the use of small particle size inert diluent lodged among the catalyst particles helps in narrowing the liquid flow channels and improves the liquid holdup. Axial dispersion is substantially reduced to the extent where plug-flow can be established, and much higher liquid hold-up improves catalyst wetting and contact efficiency. In the present work, we utilized an inert with smaller particle diameter than the commercial catalyst, and we have found that the use of this inert material in a 50/ 50 vol % ratio with the catalyst improves straight-run gas oil hydrodesulfurization conversions. Literature reports1,3 that the diluent should be even smaller to achieve better conversion. The size of diluent will depend on several factors such as size and shape of the commercial catalyst, the diameter and length of the reactor, operating conditions, conversion levels, mode of operation, type of reaction system, gas and liquid velocities, type and shape of diluent, and the technique for loading the catalyst and diluent. Apparent Activation Energies. For selecting and evaluating a new catalyst, detailed kinetic expressions for hydrodesulfurization reaction are overinformative, and catalyst activity can be simply quantified by a reaction order and an activation energy.11

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Castan˜ eda-Lo´ pez et al. Table 4. Apparent Activation Energies (in kJ/mol) for Different Systems without diluent minimun amount of catalyst (14 mL) maximum amount of catalyst (29 mL) with diluentb 50/50 vol % catalyst-to-diluent ratio 25/75 vol % catalyst-to-diluent ratio a

-1 Figure 7. S-1 vs 1/WHSV plots for different systems p - Sf at 613 K. (a) Downflow without diluent, (b) upflow without diluent, (c) upflow with diluent. (O) Minimum amount of catalyst, (b) maximum amount of catalyst, (0) 50/50 vol % catalyst-to-diluent ratio, (9) 25/75 vol % catalyst-to-diluent ratio.

For this reason we evaluated the apparent kinetic constants using the following second-order rate equation: -1 S-1 p - Sf ) kap

1 (WHSV )

downflow

123.03 129.60a

125.15 144.14

129.07 96.10

Amount of catalyst: 24 mL. b Amount of catalyst: 29 mL.

only the upflow mode of operation was utilized. It is observed that the amount of diluent in the catalytic bed has a notable effect on catalyst behavior. Apparent kinetic constants were determined with the slope of the lines presented in Figure 7, which was prepared with information generated at 613 K. Other kinetic constants were evaluated with those slopes obtained at 633 and 653 K of reaction temperatures. Apparent activation energies were calculated by using Arrhenius equation, and values are given in Table 4. The difference between activation energy values obtained without diluent for the maximun and minimun amounts of catalyst was high in the case of the downflow system, 125.15 and 144.14 kJ/mol, respectively, and it was small for the upflow system, 123.03 and 129.60 kJ/ mol, respectively. The same behavior showed the individual values of apparent kinetic constants (kap) since, for instance, at temperature of 613 K they were 5.87 and 5.97 wt %-1 h-1 for minimum and maximum amounts of catalyst using upflow system, and for downflow the difference between these kap values was higher. This also confirms the better performance of upflow mode of operation discussed in previous sections. When 25/75 vol % catalyst-to-diluent ratio is used, a low activation energy is obtained, in fact, this is the lowest value reported in Table 4. This low activation energy value is mainly due to the insufficient contact between feedstock and catalyst. Conclusions

(1)

where Sp is the sulfur content in product (wt %), Sf the sulfur content in feed (wt %), WHSV the weight-hourly space velocity, and kap the apparent kinetic constant (wt %-1 h-1). -1 Figures 7, parts a and b, show the plot of S-1 p - Sf vs 1/WHSV for undiluted beds using upflow and downflow modes of operation. It can be seen that the reduction of catalyst amount loaded to the reactor by a factor of about 2, and hence the bed length, is accompained by a change of the apparent catalyst performance. This difference in catalyst behavior is higher in downflow compared to upflow system, which is due to the reasons mentioned in the previous discussion. These findings agree very well with others reported in the literature.4 On the other hand, Figure 7c presents a comparison -1 of the plot of S-1 p - Sf vs 1/WHSV for 27/75 vol % and 50/50 vol % catalyst-to-diluent ratios. In these cases,

upflow

The upflow mode of operation is better than the downflow mode and all types of maldistribution can be avoided by feeding gas and liquid cocurrently upward through the catalyst bed. HDS conversions are the same for upflow and downflow modes when the reaction is studied at lower LHSV and higher temperature. The maldistribution of liquid can be minimized by using the proper size of diluent. Diluent can increase the liquid hold-up and hence increases HDS activity. Proper ratio of catalyst to diluent is also an important factor and it is found that a 50/50 vol % ratio shows better HDS performance in our present investigation. It was found that the best system to minimize external gradients, poor wetting of catalyst, axial dispersion, wall effects, and flow maldistribution, is upflow with a 50/50 vol % catalyst-to-diluent ratio. Acknowledgment. The authors thank Instituto Mexicano del Petro´leo for its financial support. EF010006P