H2S Ratio on the

A Pilot-Plant Simulation of the Influence of the H2/H2S Ratio on the Efficiency of a ... Kinetic parameter estimation and simulation of trickle-bed re...
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Ind. Eng. Chem. Res. 2006, 45, 7393-7398

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KINETICS, CATALYSIS, AND REACTION ENGINEERING A Pilot-Plant Simulation of the Influence of the H2/H2S Ratio on the Efficiency of a Commercial Hydrodesulfurization Plant Radoslav Micic† and Goran Boskovic*,‡ NIS-Oil Rafinery NoVi Sad, Put Sajkaskog odreda bb, and Faculty of Technology, UniVersity of NoVi Sad, 21000 NoVi Sad, Serbia

Due to a low content of H2 in the makeup gas which was added into the recirculation stream, the hydrodesulfurization (HDS) efficiency of the commercial unit is continuously decreasing. To maintain the required level of sulfur in the products, both elevated temperature and a space velocity decrease were applied. This situation ends with a severe capacity reduction. The observed decline in HDS efficiency is due to the disturbed H2/H2S ratio in the recycle gas. This was confirmed by the pilot-plant studies of the influence of the reaction parameters on the HDS efficiency as a function of H2 purity in the gas phase. 1. Introduction Hydrodesulfurization (HDS) of middle distillates is, nowadays, an unavoidable process in oil refining and one of the most broad catalytic processes in general. Its enormous importance is imposed on by the ever-tightening environmental legislation related to sulfur levels in transportation fuels, coupled with a continuously decreasing quality of oil which is offered on the world market.1,2 The worldwide standard of a 500 ppm sulfur limit in diesel has been changed recently by the new requirement for 10-15 ppm.3,4 These ultralow sulfur diesel specifications will require refineries to significantly increase their capacities. Modern trends in deep HDS are related to new catalyst systems combined with plant modification. Mainly, a catalyst charge comprises both the traditional Mo-based promoted catalyst and a noble metal-based catalyst. The latter is usually placed into an additional bed of the present reactor, or even an extra reactor may be added to the system.5 In addition, these new systems must include a recycle gas purification unit (PSA) to deliver a high partial pressure of hydrogen. The revamping of an existing plant is sometimes even impossible due to its limitation related to the maximal designed pressure. Faced with this challenge, an engineer can go close to new diesel standards only by intensifying the process parameters, mainly by decreasing the space velocity and increasing the temperature. Even then the edge of 50 ppm S obtained in the product is hard to move over. The drawbacks of these extreme process parameters are the shortening of the catalyst lifetime, more energy consumption per ton of products, and lover feed valorization. The intensification of the process parameters, which is in line with deep HDS, often causes catalyst deactivation of various mechanisms.6 Poisoning by H2S,7 loss of sulfur atoms from the active phase,8 inhibition by water,9 coking and metal deposition,10,11 and sintering9,12 have all been extensively investigated, mainly on the laboratory scale. Deactivation in a real process * To whom correspondence should be addressed. E-mail: [email protected]. Tel: 381 21 458 3629. Fax: 381 21 450 413. † NIS-Oil Rafinery Novi Sad. ‡ Faculty of Technology.

plant is not one-sided and determined by a single mechanism, however, but is rather a combination of several mechanisms as was shown in our previous work.9 Due to parameter severity for deep HDS and impurities in the low-quality feed, catalyst deactivation by the majority of the mentioned mechanisms is more or less an accepted fact for refineries. The only thing an engineer can do is to cope with the deactivation caused by an excess of self-produced H2S, keeping the balance between H2S and H2.7 In this work HDS of a diesel feed was performed in a commercial unit under various conditions to get the required S level of 50 ppm. The plant is a typical HDS unit originally designed for middle distillate fraction, with a maximal performance of 350 ppm S in the products at 54 bar and a reaction temperature of 330 °C. The commercial catalyst applied in these investigations was of the Co-Mo-based type, and the feed was with a total S content of 5400 ppm of unknown origin. To better understand the results obtained from the commercial unit, they are compared and discussed with the HDS results of a model-system feed obtained from a bench-scale laboratory reactor using the same commercial catalyst. The experimental protocols were designed to be as similar as possible to the standard working conditions used in the commercial plant. 2. Experimental Section 2.1. Commercial Plant. The semiregenerative HDS commercial plant comprises an adiabatic down-flow reactor with a distributor for directing the liquid feed. The reactor is “densely loaded” with approximately 12 m3 of Co-Mo-based commercial catalyst, in the form of extrudates 1/16 in., with a bulk density of 830 kg/m3. The designed purity of the makeup gas is 83 vol % H2, and the minimal requested hydrogen purity in the recycle gas is 70 vol %, giving 350 ppm S in the products as the maximal efficiency. The designed ratio H2/feed is 286 mN3/m3, but the real hydrogen consumption varies depending on the quality of the feed. The plant is designed for the treatment of a feed with a broad distillate range, from the kerosene to the VGO case, and with a predicted LHSV of 2.8 m3/m3. The maximal absolute pressure is 55 bar.

10.1021/ie060357x CCC: $33.50 © 2006 American Chemical Society Published on Web 09/27/2006

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Ind. Eng. Chem. Res., Vol. 45, No. 22, 2006

Figure 1. Bench-scale laboratory reactor (1, feed tank; 2, feed measuring; 3, feed piston pump; 4, gas tank; 5, preheater; 6, reactor; 7, 8, water coolers; 9, gas-liqud separator).

In a particular case a broad oil fraction (ASTM bp 180-360 °C) with 5400 ppm sulfur was HDS treated by varying the LHSV from 2.2 to 3.0 (m3‚h-1)/m3. The reactor temperature was in the range 330-360 °C, and the constant pressure was 54 bar. A flow of 5200 mN3/h of recycle gas was used, continuously enriched by 2000 mN3 of the makeup gas. The catalyst was activated by a standard protocol using an appropriate mercaptane. The catalyst performances followed in this work correspond to the early stage of its lifetime. Therefore, we assume that no substantial deactivation of the catalyst caused by metal deposition, coking, or sintering occurs during the study. The plant, the performances of which were followed in these investigations for several days, was finally stopped due to the low performances of the re-forming unit which was used for the supply of the makeup gas. 2.2. Bench-Scale Reactor. The laboratory reactor used in this investigation (Andreas Hofer, GmBH) is a down-flow stainless steal fixed-bed reactor 30 mm in inner diameter and 1000 mm long. The reactor is positioned in an integrated electrical furnace controlled by two Ni-Cr-Ni thermocouples measuring temperatures at the bottom and the top of the catalyst bed. The effective volume of the catalyst is limited to 150 cm3. A liquid feed, pumped by a high-pressure piston pump of maximal capacity 2.0 L/h, is first preheated and mixed with hydrogen at the reactor entrance (Figure 1). Leaving the reactor, the products are first cooled in two water coolers and then directed to a vapor-liquid separator. The pressure in the system is measured by a manometer at the top of the reactor and controlled by a valve positioned at the top of the separator. Liquid products are collected at the bottom of the separator from which samples for analysis of the S amount are taken periodically. In total, eight experiments were performed covering several reaction temperatures with constant LHSV and several flow rates at a particular temperature. The altitudes of these parameters

were chosen in a way to get a demanded S level in the product with the simultaneous influence of the gas-phase purity. To reach a quasi-steady state, every experiment lasted for 2 h before the sample was taken for S analysis. Light gas oil (LGO) was chosen as a test feed, with an average S content of 4200 ppm. The feed rate was varied from 160 to 220 cm3/h, resulting in space velocities from 2.19 to 2.92 h-1. The gas flow of N2 and H2, premixed in various ratios, was kept constant at 40400 cm3/h to satisfy the requested H2/feed ratio of 120-200. The experiments were performed as a one-through gas stream. The reaction temperature was chosen in the range of 330-360 °C. The same commercial catalyst used in the commercial plant was loaded into the bench reactor to exclude diffusion effects. The effective catalyst particle diameter is calculated as dp ) 0.82 mm.13 The ratio between the bed length and the effective particle diameter, L/dp ) 305, was calculated assuming a firstorder reaction and a very high conversion of 97%. Since this ratio does not satisfy the required minimal value for a threephase system,14 the catalyst bed had to be diluted with an inert material. Therefore, crushed R-alumina particles of size less than 0.2 mm were mixed with the catalyst in the ratio 1/1, resulting in a total volume of the catalyst and inert of approximately 150 cm3. The axial diffusion criterion is calculated by the ratio of the inner reactor diameter and the effective particle diameter as D/dp ) 36.5. It proves that there is no implication of axial dispersion since the value is higher than 25, which is the limit required for the presence of a three-phase system.14 The wetting criterion was checked by means of the equation

W)

ηLuL > 5 × 10-6 FLdpg

(1)

where ηL, FL, and uL are the dynamic viscosity, density, and linear velocity, all related to the liquid phase, g is the gravity constant, and W is a dimensionless number comparing flowing friction forces and the gravity force. Good catalyst bed wetting with an appropriate reconstruction of the liquid film at the top of the catalyst surface is guaranteed by the range of the calculated wetting criterion as W ) (1.9-2.6) × 10-5 obtained for the extremes of linear velocity.14 The catalyst was activated in a flow of H2S/H2 mixture very rich in H2. The amount of S added was calculated to be 3-4 wt %, the same as in the commercial unit. Activation was performed with gradually increased temperature up to 320 °C, which lasted for 1 h. 2.3. Analysis. Sulfur in the feed and product in both the commercial and pilot-plant studies was analyzed by the EDXR method on LABX-3500, Oxford Instruments, analytical equipment. The compositions of the gas streams from the commercial unit were analyzed in the following way: H2S in the recycle gas was analyzed by the Tutwiler-UOP-9 titration method. The gas chromatography UOP-539 method applied to gas chromatograph HP 5880 and a combination of Sebaconitryl, Porapac-Q, and molecular sieves 5A columns was used for the analysis of the compositions of both the recycle and makeup gases. The composition of the one-through gas stream from the bench reactor was not measured, but the resulting H2/H2S ratio was calculated on the basis of the hydrogen in the inlet stream. The partial pressure of H2S was calculated using the formula

pH2S )

Vm p F (4200 ppm)X MS H2/feed ratio feed

(2)

Ind. Eng. Chem. Res., Vol. 45, No. 22, 2006 7395 Table 1. Reaction Conditions and Results of HDS on the Commercial Unit

time-on-stream, h makeup gas composition, % 0 4 8 12 16 makeup gas composition, % 20 24 28 32 36 makeup gas composition, % 40 44 makeup gas composition, % 48 52 56 60 makeup gas composition, % 64 68 makeup gas composition, % 72 76 80 84 makeup gas composition, % 88 92 96

T, °C

S level in product, ppm

recycle gas composition, % capacity, m3/h

H2S

330 340 345 350 360

420 260 220 150 120

37.8 37.8 37.8 37.8 37.8

1.4 1.4 1.4 1.5 1.2

330 340 345 350 360

490 320 190 110 60

35.1 35.1 35.1 35.1 35.1

1.3 1.6 1.5 1.4 1.5

330 340

430 270

31.5 31.5

1.3 1.4

345 350 360 360

300 340 130 100

31.5 31.5 31.5 28.1

1.3 1.4 1.4 1.3

350 345

110 170

28.1 28.1

1.6 1.6

340 330 330 340

210 270 340 230

28.1 28.1 26.3 26.3

1.4 1.6 1.3 1.3

345 350 360

170 180 80

26.3 26.3 26.3

1.3 1.3 1.3

where MS is the molar mass of sulfur, Vm molar volume at standard conditions, p is total pressure, F is liquid feed density (840 kg/m3), and X is conversion.7 3. Results and Discussion The reaction conditions and HDS plant results with time-onstream are presented in Table 1. At the top of the table the composition of the makeup gas is presented, followed by the composition of the recycle gas, which was measured in constant time intervals of 4 h. The makeup gas analyses were periodically repeated after a couple of hours of time-on-stream, and these results are also presented in a real-time experiment in Table 1. The task of the makeup stream coming from the re-forming unit is to refresh the recycle gas responsible for HDS. After their mixing, due to different gas-liquid equilibrium dependencies on pressure for different constituents, a gas-phase

Figure 2. Change of the H2S and H2 fractions in the recycle gas of the commercial unit with time-on-stream.

H2

C1

C2

C3

C4

76.6 80.8 80.8 79.9 79.9 80.1 76.2 79.6 78.4 78.7 79.7 79.5 74.6 80.0 78.6 73.7 78.1 77.6 76.7 76.3 72.6 76.0 76.5 72.3 76.4 75.4 75.7 75.7 72.8 74.6 74.5 74.6

15.1 15.9 16.0 16.8 16.8 16.8 15.2 17.1 17.9 17.5 17.0 17.1 15.8 17.0 18.1 16.4 18.6 19.0 19.6 20.2 17.0 20.2 19.8 17.0 20.0 20.9 20.9 20.9 16.7 21.8 21.9 21.8

5.6 1.0 1.0 1.2 1.0 1.1 5.7 1.2 1.3 1.3 1.1 1.1 6.3 1.1 1.1 6.5 1.2 1.2 1.4 1.4 6.7 1.4 1.3 6.7 1.3 1.4 1.3 1.3 6.6 1.5 1.5 1.5

1.4 0.1 0.1 0.1 0.1 0.1 1.5 0.1 0.1 0.2 0.1 0.1 1.8 0.1 0.1 1.8 0.1 0.1 0.1 0.1 1.9 0.1 0.1 1.9 0.1 0.1 0.1 0.1 2.0 0.2 0.1 0.2

0.7 0.8 0.7 0.7 0.7 0.7 0.8 0.7 0.8 0.8 0.6 0.7 0.9 0.6 0.6 0.9 0.8 0.6 0.7 0.7 1.0 0.7 0.7 1.0 0.7 0.6 0.6 0.6 1.1 0.7 0.6 0.6

deficiency in hydrocarbons occurs. This leads to H2 fraction enrichment in the recycle gas, being crucial for the HDS process. A tendency of decreasing fraction of hydrogen in the recycle gas can be noticed from Table 1 with time-on-stream, and this behavior is presented in Figure 2. Considering the relatively low amount of H2 spent for the S removal alone, this lack of H2 is probably due to a permanent decrease of the quality of the makeup gas. At higher temperatures the contribution of cracking and hydrogenation reactions to the hydrogen depletion of the recycle gas may be considered. However, this input is small compared to the main reason, which is the reduction of the quality of the makeup gas. A trend opposite that manifested by H2 with time-on-stream is to be expected for H2S as well. However, due to its low concentration in the recycle gas and an inappropriate analytical method used for its determination, this was not observed (Figure 2). To meet as low a S content in the products as possible, the reaction temperature was gradually increased, resulting in an expected trend of S removal. In that way, the level of sulfur in the products can be decreased maximally to 120 ppm. To push the S level lower than 100 ppm, however, the reaction temperature increase had to be coupled with the lowering of the liquid feed capacity (Table 1). The effect of the LHSV reduction on the S level in the products is not proportional, however, but decreases due to the continuous H2 fraction reduction. The HDS efficiency as a function of both the feed capacity and H2 fraction in the recycle gas is presented in Figure 3 for the highest applied temperature of 360 °C obtained at different time-on-stream points. It can be seen that the capacity decrease from 35 to 31 m3/h brings an unexpected S level increase from 60 to 130 ppm due to a dramatic decrease of the H2 fraction in the gas phase as high as 3%. A further decrease in the H2 purity with time-on-stream is not so distinct, however,

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Ind. Eng. Chem. Res., Vol. 45, No. 22, 2006 Table 2. Reaction Temperature Necessary To Obtain a S Level of 150 ppm in the Product for a Given H2 Fraction in the Gas Phasea T, °C

gas flow, cm3/h

H2 fraction, %

H2 flow, cm3/h

H2/feed ratio, vol/vol

H2/H2S ratio, vol/vol

H 2S pressure, bar

330 336 342 360

44000 44000 44000 44000

100 82 72 60

44000 36080 31680 26400

200 164 144 120

83.98 68.87 60.47 50.39

0.70 0.86 0.97 1.17

a

Reaction conditions: laboratory reactor, 55 bar, LHSV ) 2.9 h-1.

Figure 3. Coupled effect of the feed capacity and recycle gas purity on the S level in the product obtained at 360 °C at different times-on-stream of the commercial process.

and can be compensated by the capacity decrease, resulting in the accepted level of S of 80 ppm (Figure 3). The negative influence of H2S on the HDS efficiency of both commercial and model-system catalysts has been investigated in the past.15,16 In the case of the HDS of LGO over commercial Co-Mo/Al2O3, a substantial retardation of the HDS of dibenzothiophenes (DBTs) and the corresponding monosubstituted (4-MDBTs) and disubstituted (4,6-DMBTs) dibenzothiophenes was found when a real gas-to-oil ratio of 125 vol/vol was applied. It was attributed to the poisoning effect of the solely self-produced H2S, since the addition of as much as 3% H2S to the feed was found to be without any negative effect on the HDS of 4,6-DMBTs.15 In the case of a bulk MoS2 model-system HDS catalyst processing a feed of 1% DBT in decane, a decrease of the biphenyl (BP) product was explained by the self-produced H2S and corresponding self-poisoning of the direct desulfurization route (DDS) responsible for the direct S-C cleavage.16 The results of the HDS of LGO having a high S content (8445 ppm) have been reported recently, pointing out the importance of the H2 partial pressure.7 Namely, inhibition by the selfproduced H2S plays an important role in HDS, which occurs in two parallel directions: a DDS and a hydrogenation route. When no H2S is present, the partial pressure of H2 is not of crucial importance for the desulfurization route occurring on Svacancies. It is important only for the hydrogenation route on SH-sites. However, when a partial pressure of H2S as low as 0.2 bar is present, the amount of H2 becomes very important for the creation of new vacancies for the direct desulfurization. Since this route is much faster than the previous one, the resulting HDS activity is low.17 Therefore, the H2/H2S ratio in the gas phase alters the balance between different active sites and the corresponding HDS routes and reflects the HDS total activity.7 The results obtained from the present investigation on the commercial plant show temperature as an important factor in determining the HDS efficiency. The data, however, indicate that the quality of the recycle gas is of high importance as well. Once the H2 content in the recycle gas is low, a substantial S removal can be obtained only by combining the elevated reaction temperature with the feed capacity decrease. This low H2 content brings an unsuitable H2/H2S ratio, resulting in the H2S poisoning of S-vacancies and the inhibition of the direct desulfurization route. Consequently, the HDS activity declines. Therefore, the quality of the recirculation gas is as important a parameter as

Figure 4. Interdependence of temperature and H2/H2S ratio at LHSV ) 2.9 h-1 to obtain a constant S level in the product of 150 ppm.

the reaction temperature. Which one of these two factors is more important is not possible to estimate from the data in Table 1, however. On the basis of the previous results, two series of independent experiments were performed in the laboratory on the benchscale reactor. The aim was to find the functional dependence of the outlet S on the reaction temperature and space velocity, with the simultaneous influence of the gas-phase purity. Experiments were designed to meet the best performances of the catalyst obtained in the commercial plant, and only results meeting these conditions are presented. The process parameters of the first set of experiments, giving products with a constant sulfur level of 150 ppm, are presented in Table 2. The reaction temperature was changed in the range 330-360 °C and the feed rate kept constant at 220 cm3/h, resulting in LHSV ) 2.9 h-1. The absolute pressure was 55 bar in the course of the whole experiment. A constant gas flow of 44 L/h was applied, with a variable H2 fraction from 60% to 100%, the rest being N2. The magnitude of H2S in the outlet gas phase calculated as the difference between S in the feed and that in the product amounted to 0.52 L/h, as shown in Table 2. The results presented in Table 2 undoubtedly show that a temperature increase is not enough to reach deep HDS. The positive influence of 30 °C is nullified by the decrease of the H2/H2S ratio. In fact, in reality, this ratio is even lower due to the fact that the largest amount of hydrogen is consumed in reactions other than HDS, such as hydrogenation, cracking, etc. As seen from Figure 4, as long as the H2/H2S ratio is high, a relatively low positive difference in temperature is enough to reach the required S level of 150 ppm. With a H2 fraction of 60% in the inlet gas, which corresponds to a partial pressure of H2S close to 1 bar (Table 2), a significant temperature increase of 18 °C is necessary to keep the same level of HDS efficiency (Figure 4). The results of the pilot-plant experiments reported earlier18 are quite similar to the results of this investigation, pointing out, however, that temperature is a major factor in determining the HDS efficiency. A substantial increase of the contact time was necessary to reach a S level of up to 300 ppm but not before the maximal reaction temperature of 370 °C was applied.18 This

Ind. Eng. Chem. Res., Vol. 45, No. 22, 2006 7397 Table 3. Space Velocity Necessary To Obtain a S Level of 50 ppm in the Product for a Given H2 Fraction in the Gas Phasea LHSV, h-1

gas flow, cm3/h

H2 fraction, %

H2 flow, cm3/h

feed rate, cm3/h

H2/feed ratio, vol/vol

H2S amt, L/h

H2/H2S ratio, vol/vol

H2S pressure, bar

2.9 2.8 2.7 2.1

44000 44000 44000 44000

100 88 79 60

44000 34760 27720 17600

220 210 200 160

200 184.38 173.8 165

0.54 0.51 0.49 0.39

81.96 75.56 71.22 67.62

0.72 0.78 0.83 0.87

a

Reaction conditions: laboratory reactor, 55 bar, T ) 360 °C.

82 vol/vol, it was possible to obtain 50 ppm S in the product (Table 3). At the same time, using the same reaction temperature and LHSV, a product with a S level of 150 ppm was reached in the first set of experiments (Table 2), due to a very low H2/H2S ratio, such as 50 vol/vol. 4. Conclusions

Figure 5. Interdependence of LHSV and H2/H2S ratio at 360 °C to obtain a constant S level in the products of 50 ppm.

much lower HDS efficiency in comparison to the results of the present study may be attributed to the difference in the S level in the feed (almost 2 wt % S vs 0.4 wt % S in the present investigation), as well as the different pressure and higher feed end point, indicating a higher content of highly unreactive disubstituted dibenzothiophenes.17 The second set of experiments in the pilot plant comprises the space velocity change from 2.1 to 2.9 h-1, while the reaction temperature and gas flow were kept constant at 360 °C and 44 L/h, respectively (Table 3). The variable H2 fraction in the gas flow was applied from 60% to 100%, the rest being N2. The absolute pressure was constant at 55 bar. These parameters gave products with a constant sulfur content of 50 ppm, which is the best result obtained in both the pilot-plant and commercial investigations. The amount of H2S released and the H2/H2S ratio were calculated in the same way as in the previous set of experiments. The results obtained, based on the experiments given in Table 3, represent a complex function of different parameters. On one side, it is obvious that the decrease of the space velocity at a constant reaction temperature corresponds to the efficiency of the S removal, as shown in Figure 5. It is very effective for high H2/H2S ratios having an almost linear dependence on the requested 50 ppm S, while for lower values of H2/H2S ratios a considerable decrease of LHSV is necessary. The influence of the LHSV decrease is dual, however. First, it makes it possible to obtain a very low level of S at 50 ppm in the products by keeping the production of H2S lower due to a high H2/feed ratio. As a result, the H2/H2S proportion is higher compared to the data in Table 2, causing lower poisoning of S-vacancies responsible for the direct desulfurization route.7 Second, by increasing the contact time, the slower hydrogenation route is emphasized, although the price is paid by a lower capacity. The outcome of both effects is a considerable HDS activity. The results presented in Figures 4 and 5 clearly show the dominant influence of the gas purity over the amount of H2S self-produced. Namely, the amount of self-produced H2S at the highest LHSV applied in the second set of experiments (Figure 5) is higher in comparison to the corresponding H2S level produced at 360 °C in the first set of experiments (Figure 4). However, due to a higher corresponding H2/H2S ratio, such as

By using a classical Co-Mo-based commercial catalyst in a standard HDS plant, diesel products with a level of sulfur up to 50 ppm can be obtained. This sulfur level presently satisfies standards in East European countries in transition. To maintain this requirement, an increase of the reaction temperature must be coupled with a space velocity decrease. However, the recycle gas purity must be considered as well. A gas rich in selfproduced H2S poisons the S-vacancies responsible for the direct desulfurization route, resulting in a substantial decrease of the total HDS activity. By keeping the H2/H2S ratio high enough, this situation can be prevented. Acknowledgment The financial support of the Serbian Ministry of Science and Environmental Protection (Project No. 142024) is highly appreciated. Literature Cited (1) Fujikawa, T.; Kimura, H.; Kiriyama, K.; Hagiwara, K. Development of ultra-deep HDS catalyst for production of clean diesel fuels. Catal. Today 2006, 111, 188. (2) Topsoe, H.; Clausen, B. S.; Massoth, F. E. Hydrotreating Catalysts. In Catalysis: Science and Technology; Anderson, J. R., Boudart, M., Eds.; Springer: Berlin, 1996; Vol. 11, p 1. (3) Song, C. An Overview of New Approaches to Deep Desulfurization for Ultra-Clean Gasoline, Diesel Fuel and Jet Fuel. Catal. Today 2003, 86, 211. (4) Segawa, K.; Takahashi, K.; Satoh, S. Development of new catalysts for deep hydrodesulphurization of gas oil. Catal. Today 2000, 63, 123. (5) Bej, S. K. Revamping of diesel hydrodesulfurizers: options available and future research needs. Fuel Process. Technol. 2004, 85, 1503. (6) Furimsky, E.; Massoth, F. E. Deactivation of hydroprocessing catalysts. Catal. Today 1999, 52, 381. (7) Vogelaar, B. M.; Kagami, N.; van Langeveld, A. D.; Eijsbouts, S.; Moulijn, J. A. Active sites and activity in HDS catalysis: the effect of H2 and H2S partial pressure. Prepr. Pap.sAm. Chem. Soc., DiV. Fuel Chem. 2003, 48, 548. (8) Vogelaar, B. M.; Steiner, P.; van Langeveld, A. D.; Eijsbouts, S.; Moulijn, J. A. Deactivation of Mo/Al2O3 and NiMo/Al2O3 catalysts during hydrodesulphurization of thiophene. Appl. Catal., A 2003, 251, 85. (9) Neducin, R. M.; Boskovic, G.; Kis, E.; Lomic, G.; Hantsche, H.; Micic, R.; Pavlovic, P. Deactivation of industrial hydrotreating catalyst for middle petroleum fractions processing. Appl. Catal., A 1994, 107, 133. (10) Gosselink, J. W.; van Veen, J. A. R. Coping with catalyst deactivation in hydrocarbon processing. In Catalyst DeactiVation 1999; Delmon, B., Froment, G. F., Eds. Stud. Surf. Sci. Catal. 1999, 126, 3. (11) Seki, H.; Yoshimoto, M. Deactivation of HDS catalyst in two-stage RDS process: II. Effect of crude oil and deactivation mechanism. Fuel Process. Technol. 2001, 69, 229. (12) Topsoe, H.; Clausen, B. S.; Massoth, F. E. Hydrotreating Catalysts. In Catalysis: Science and Technology; Anderson, J. R., Boudart, M., Eds.; Springer: Berlin, 1996; Vol. 11, p 1.

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(13) Sakiadis, B. C. Fluid and particle mechanics. In Perry’s Chemical Engineers’ Handbook, 6th ed.; Perry, R. H., Green, D. W., Maloney, J. O., Eds.; McGraw-Hill: New York, 1984. (14) Sie, S. T. Advantages, possibilities and limitations of small-scale testing of catalysts for fixed-bed processes. In DeactiVation and Testing of Hydrocarbon-Processing Catalysts; O’Connor, P., Takatsuka, T., Woolery, G. L., Eds.; ACS Symposium Series 634, 6; American Chemical Society: Washington, DC, 1996. (15) Kabe, T.; Akamatsu, K.; Ishihara, A.; Otsuki, S.; Godo, M.; Zhang, Q.; Qian, W. Deep Hydrodesulfurization of Light Gas Oil. 1. Kinetics and Mechanisms of Dibenzothiophene Hydrodesulfurization. Ind. Eng. Chem. Res. 1997, 36, 5146. (16) Farag, H.; Sakanishi, K.; Kouzu, M.; Matsumura, A.; Sugimoto, Y.; Saito, Investigation of the Influence of H2S on Hydrodesulfurization

of Dibenzothiophene over a Bulk MoS2 Catalyst. Ind. Eng. Chem. Res. 2003, 42, 306. (17) Hu, M. C.; Ring, Z.; Briker, J.; Te, M. Rigorous hydrotreater simulation. Pet. Technol. Q. 2002, Spring, 85. (18) Lappas, A. A.; Budisteanu, R.; Drakaki, K.; Vasalos, I. A. Production of low aromatics and low sulfur diesel in a hydrodesulphurization (HDS) pilot plan unit. Global Nest 1999, 1, 15.

ReceiVed for reView March 23, 2006 ReVised manuscript receiVed June 9, 2006 Accepted August 15, 2006 IE060357X