Comparison of Quench Systems in Commercial Fixed-Bed

Energy & Fuels 2007, 21, 1133-1144

1133

Comparison of Quench Systems in Commercial Fixed-Bed Hydroprocessing Reactors Anton Alvarez,† Jorge Ancheyta,*,†,‡ and Jose´ A. D. Munõz† Instituto Mexicano del Petro´ leo, Eje Central La´ zaro Ca´ rdenas 152, Col. San Bartolo Atepehuacan, Me´ xico D.F. 07730, and Escuela Superior de Ingenierıá Quı´mica e Industrias ExtractiVas (ESIQIE-IPN), Me´ xico D.F. 07738 ReceiVed October 16, 2006. ReVised Manuscript ReceiVed December 22, 2006

Several aspects of quenching in fixed-bed hydroprocessing reactors have been reviewed. Examples of processes using hydrogen and liquids as quench fluids are described, and the effects of each quench fluid on the reaction system are discussed. The advantages and disadvantages of each approach on the product quality and process configuration are also highlighted. It was recognized that traditional quench hydrogen helps in improving product quality and reducing catalyst deactivation by coke formation but increases equipment requirements (e.g., compressor) and relative costs, while quench liquid avoids this problem but increases reactor volume.

1. Introduction Catalytic hydroprocessing is a mature technology practiced in the petroleum refining industry for upgrading hydrocarbon streams for the last 60 years.1 Hydroprocessing can be divided in two main types of processes: hydrotreatment (HDT) and hydrocracking (HYC).2 During HDT, when processing light and middle distillates, reactions such as hydrodesulfurization (HDS), hydrodenitrogenation (HDN), and hydrodearomatization (HDA) occur, bringing down the content of those impurities substantially; when handling heavier feeds, reactions such as hydrodemetallization (HDM) and hydrodeasphaltenization (HDA) are also present.3 Hydrocracking is the other type of process employed for converting various hydrocarbons of higher boiling point ranges into more valuable products such as gasoline, diesel, and jet fuel.4 Numerous hydroprocessing technologies are commercially available for handling all types of refinery streams.5 They differ mainly in the reactor technology, catalyst type, and composition, operating conditions and process configuration. A comprehensive review dealing with various aspects of hydroprocessing of heavy oils has been reported recently.6 Operating conditions vary depending on the type of feedstock and the desired quality of the products; typical industrial hydroprocessing units operate at high pressures (2-20 MPa) and temperatures (320-440 °C), and the hydrogen-to-oil (H2/ oil) ratio and space velocity (LHSV) vary from 2000 to 10 000 SCF/bbl and from 0.2 to 8 h-1, respectively.7,8,9 * Corresponding author. Fax: (01-55) 9175-8429. E-mail: [email protected]. † Instituto Mexicano del Petro ´ leo. ‡ Escuela Superior de Ingenierıá Quı´mica e Industrias Extractivas (ESIQIE-IPN). (1) Mederos, F. S.; Rodrı´guez, M. A.; Ancheyta, J.; Arce, E. Energy Fuels 2006, 20, 936-945. (2) Speight, J. Catal. Today 2004, 98, 55-60. (3) Ancheyta, J.; Betancourt, G.; Marroquıń, G.; Centeno, G.; Castan˜eda, L. C.; Alonso, F.; Munõz, J. A.; Go´mez, M. T.; Rayo, P. Appl. Catal. A 2002, 233, 159-170. (4) Mohanty, S.; Kunzru, D.; Saraf, D. N. Fuel 1990, 69, 1467-1473. (5) Rana, M. S.; Sa´mano, V.; Ancheyta, J.; Diaz, J. A. I. Fuel 2007, 86, 1216-1231. (6) Ancheyta, J.; Rana, M. S.; Furimsky, E. Catal. Today 2005, 105, 3-15. (7) Satterfield, C. N. AIChE J. 1975, 21, 209.

The selection of a hydroprocessing reactor technology is directly influenced by the quality of the feed (e.g., metals and asphaltenes contents), in other words, by the catalyst deactivation rate.10 Traditionally, fixed-bed reactors were employed for handling light feeds, but they were gradually adapted for processing heavier feeds such as vacuum gasoil and atmospheric/ vacuum residua.11-13 Examples of these processes are the HYVAHL process developed by IFP14 and UOP’s Unibon BOC and RCD processes.15 The main disadvantage of fixed-bed reactors when processing heavy feeds is the premature catalyst deactivation, which leads to loss in catalyst activity; during the time-on-stream, this reduces drastically the length of the run due to the frequent shutdowns required for replacing the catalyst.10,16,17 Catalyst replacement is no longer an issue in ebullated-bed and moving-bed reactors where the main feature is the capability of the process to replace spent catalyst without interrupting the operation.10,11 However, the problem turns now into one of sediment formation, which is actually an important concern during commercial operation of these reactors. In fact, for ebullated-bed based technologies, sediment formation is normally limited to values between 0.8 and 1.0 wt %. Nevertheless, among all the commercially proven reactor technologies, fixed-bed reactors are still the most used ones due to their relative simplicity and ease of operation.11 Fixed-bed reactors with various catalyst beds are frequently employed in commercial hydroprocessing. The operation of such reactors is considered to be very close to adiabatic because the (8) Bej, S. K. Energy Fuels 2002, 16, 774-784. (9) Kundu, A.; Nigam, K. D. P.; Duquenne, A. M.; Delmas, H. ReV. Chem. Eng. 2003, 19, 531-605. (10) Sie, S. T. Appl. Catal. A 2001, 212, 129-151. (11) Furimsky, E. Appl. Catal. A 1998, 171, 177-206. (12) Beaton, W. I.; Bertolacini, R. J. Catal. ReV. Sci. Eng. 1991, 33, 281-317. (13) Scheuerman, G. L.; Johnson, D. R.; Reynolds, B. E.; Bachtel, R. W.; Threlkel, R. S. Fuel Process. Technol. 1993, 35, 39-54. (14) Billon, A.; Peries, J. P.; Espeillac, M.; des Courrieres, T. Presented at the NPRA Anual Meeting, San Antonio, TX, March 17-19, 1991. (15) Gray, M. R. Upgrading Petroleum Residues and HeaVy Oils; Marcel Dekker: New York, 1994. (16) Kressmann, S.; Morel, F.; Harle´, V.; Kasztelan, S. Catal. Today 1998, 43, 203-215. (17) Chou, T. Pet. Technol. Q. 2004, 4, 79-85.

10.1021/ef060515x CCC: $37.00 © 2007 American Chemical Society Published on Web 02/07/2007

1134 Energy & Fuels, Vol. 21, No. 2, 2007

heat losses from the reactor are usually negligible compared with the heat generated by the hydroprocessing reactions.18 Catalysts with different functionalities are usually utilized to achieve the desired levels of impurities removal and hydrocracking.13 Since reaction exothermality can provoke increases in reactor temperature beyond acceptable limits, quenching along the reactor is frequently practiced. Hence, in the sense of temperature control, multiple catalytic beds, injection of quench fluids, and/or heat integration of the effluent from each bed come together.7,8,19 Additionally, in some cases, multibed systems have the provision to scrub hydrogen sulfide (H2S) and/or ammonia (NH3) from a bed effluent, which avoids inhibition of the HDS reaction in the following bed.11,20 Commonly, the total volume of catalyst is divided in two to six beds8 (although some hydrocraking reactors can have up to 30 beds),21 each one being about 3-6 m deep,7 and depending on the capacity of the plant, the diameter of the reactor can be from 1 to 6 m.22 Modern hydrotreatment reactors for middle distillates have generally two beds, and the maximum bed length may reach 12 m.8 Typical hydrocracking units employ reactors in series: one for HDT as a pretreatment stage and another one (or more) for hydrocracking;23 the reactors may have four to six beds, and the length of each bed may reach 6 m.24 Apart from the required levels of impurities in the product and feed properties, the number of catalyst beds and their lengths are determined by the temperature rise caused by the exothermality of hydroprocessing reactions. The axial temperature profile of the reactor is monitored by temperature indicators along the catalytic bed. It can be also evaluated by solving simultaneously the mass and energy balances of the reaction system.1 The bed length is limited when the reactor temperature reaches a maximum allowable temperature which is commonly about 30 °C or less above the inlet reactor temperature.7,25 Figure 1 illustrates this concept for a three-bed catalytic reactor with two quenches. The reaction temperature is an operating condition which has an enormous impact on the conversion degree of the reactants and on the catalyst cycle life, especially when hydroprocessing heavy feeds, and thus on the overall economics of the process. Increasing the average reactor temperature enhances the hydrotreating reaction rates; this approach has been practiced in order to achieve deep desulfurization of gasoil.26,27 However, temperature has a limit dictated by mechanical restrictions or by excessive hydrocracking of the feed such as in the case of gasoil, which reduces the desired product yield. On the other hand, in hydrocracking units, increasing the temperature may have a negative impact on the product distribution due to the strong influence of this process variable on hydrocracking selectivity. When the temperature exceeds 370-380 °C, the yields of methane, ethane, and propane increase considerably (18) Shah, Y. T.; Paraskos, J. A. Chem. Eng. Sci. 1975, 30, 1169-1176. (19) Chen, Y. H.; Yu, C. C. Ind. Eng. Chem. Res. 2003, 42, 27912808. (20) Sie, S. T. Fuel Proc. Technol. 1999, 61, 149-171. (21) Hsu, C. S.; Robinson, P. R. Practical AdVances in Petroleum Processing; Springer: New York, 2006; Vol. 1. (22) Speight, J. G. The Desulfurization of HeaVy Oils and Residua; Marcel Dekker: New York, 2000. (23) Bhutani, N.; Ray, A. K.; Rangaiah, G. P. Ind. Eng. Chem. Res. 2006, 45, 1354-1372. (24) Scherzer, J.; Gruia, A. J. Hydrocraking Science and Technology; Marcel Dekker: New York, 2000. (25) Zhukova, T. B.; Pisarenko, V. N.; Kafarov, V. V. Int. Chem. Eng. 1990, 30, 57-102. (26) Song, C. Catal. Today 2003, 86, 211-263. (27) Bharvani, R. R.; Henderson, R. S. Hydrocarbon Process. 2002, 81, 61-64.

AlVarez et al.

Figure 1. Three-bed reactor with two quenches.

at the expenses mainly of gasoline yield. This shift increases hydrogen consumption and reactor delta-T, while the hydrogen purity of the recycle gas decreases and the catalyst aging rate increases.28 Also, high temperatures may lead to hot spot formation which results in enhanced coke formation and catalyst sintering;29,30 such an effect reduces the catalyst cycle life, and thus, more frequent shut downs of the units are required, affecting the global profitability of the process. Therefore, temperature control in fixed-bed reactors becomes very important in order to extend the catalyst cycle life and to keep the product quality at desired levels.31 Commonly, control of the reaction temperature in hydroprocessing reactors is achieved by introducing part of the hydrogen recycle stream between the catalytic beds, so-called “quenching” or sometimes “cold shot cooling”.7,11 The use of quench liquids has been also reported. Quenching fluids are introduced in the quench zone or quench box which is typically a mixing chamber where the bed effluent is mixed with the cooling medium (Figure 2).31-33 The flow of injected fluid to each quench location is adjusted to achieve the desired temperature profile and is specified to limit the temperature rise below the maximum allowable temperature.7,25 The amount of quench fluid is calculated by solving the energy balance of the quench zone which is represented as a mixer of the quench stream and the effluent from the previous bed. Quench hydrogen, being the main reactant in hydroprocessing, has the advantage of replenishing for some of the chemically consumed hydrogen in the catalytic beds, decreasing the hydrogen sulfide and ammonia partial pressures in the reactor, which reduces the inhibition effect on HDT reactions,27,34,35 and also keeping the catalyst clean by inhibiting coke formation.21 (28) Yan, T. Y. Can. J. Chem. Eng. 1980, 58, 259-266. (29) Hanika, J.; Sporka, K.; Ruzicka, V.; Pistek, R. Chem. Eng. Sci. 1977, 32, 525-528. (30) Furimsky, E.; Massoth, F. E. Catal. Today 1999, 52, 381-495. (31) Munõz, J. A. D.; Alvarez, A.; Ancheyta, J.; Rodrı´guez, M. A.; Marroquıń, G. Catal. Today 2005, 109, 214-218. (32) Ouwerkerk, C. E. D.; Bratland, E. S.; Hagan, A. P.; Kikkert, B. L. J. P.; Zonnevylle, M. C. Pet. Technol. Q. 1999, 2, 21-30. (33) Ancheyta, J.; Speight, J. Hydroprocessing of HeaVy Oils and Residua; Taylor and Francis Group: New York, 2007. (34) Van Hasselt, B. W.; Lebens, P. J. M.; Calis, H. P. A.; Kapteijn, F.; Sie, S. T.; Moulijn, J. A.; van den Bleek, C. M. Chem. Eng. Sci. 1999, 54, 4791-4799. (35) Turner, J.; Reisdorf, M. Hydrocarbon Process. 2004, March.

Quench Systems in Fixed-Bed Hydroprocessing

Energy & Fuels, Vol. 21, No. 2, 2007 1135

Figure 2. Hydroprocessing reactor.33

The availability of quench hydrogen depends on the H2/oil ratio along the reactors, which is a design condition that influences product quality. This value of the H2/oil ratio depends primarily on the compressor capacity within the plant. High H2/oil ratios improve the product quality and increase the quench availability, for example, a staged hydrocracking unit which operates at an H2/oil ratio of ∼10 000 SCF/bbl may have up to five gas injection points. However, high H2/oil ratios also imply higher hydrogen recycling rates, and therefore, larger compressors and equipment in the separation section are required which increases the investment costs.31,36 Although hydrogen quenching is used in most of the hydroprocessing units, quench liquids sometimes are also employed, particularly cold hydrocarbon streams. Quenching with liquids may result in being advantageous due to their higher heat capacity and lower compression cost.28 Nevertheless, quench liquid may have a slight impact on the properties of the mixture and conversion degree. Technical information on this subject is really scarce and of extreme relevance for the design and optimization of hydroprocessing units. The aim of this article is then to review and discuss several aspects of temperature control in fixed-bed hydroprocessing reactors using quench fluids, focusing on the processes that use hydrogen and liquid quenching and on their effect on the reaction system behavior. (36) Medeiros, J. L.; Barbosa, L. C.; Vargas, F. M.; Arau´jo, O. Q. F.; Silva, R. M. F. Braz. J. Chem. Eng. 2004, 21, 317-324.

2. State-of-the-Art 2.1. Conventional Quenching. As said before, quench hydrogen gas is widely used in most of the hydroprocessing units. Numerous commercial processes employ this approach. Some examples of these processes are the following: (1) FCC naphtha HDS/octane recovery processes like Exxon Mobil’s OCTGain and SCAN-fining37 and the UOP-INTEVEP’s ISAL process;38 (2) gasoil HDT processes such as conventional HDS,20 Haldor-Topsøe’s staged HDS/HDA,39,40 Chevron’s vacuum gasoil HDT,41,42 and Exxon’s ULSD technologies;37,39 (3) gasoil HYC, for instance, Chevron’s Isocraking process,42,43 Exxon Mobil’s moderate pressure hydrocracking (MPHC) process,37,39 Shell’s HYC process39 and UOP’s Unibon staged HYC pro(37) ExxonMobil. Refining Technologies. http://www.prod.exxonmobil.com/refiningtechnologies (accessed Aug 2006). (38) UOP. Refining. http://www.uop.com/refining (accessed Aug 2006). (39) Refining Processes; Hydrocarbon Processing Magazine: Houston, TX, 2004. (40) Haldor Topsøe. Refining Technologies. http://www.topsoe.com (accessed Aug 2006). (41) Mukherjee, U. K.; Louie, W. S.; Dahlberg, A. J. Process for the Production of High Quality Distillates from Mild Hydrocrackers and Vacuum Gas Oil Hydrotreaters in Combination with External Feeds in the Middle Distillate Range. U.S. Patent 6,787,025, 2004. (42) Chevron. CLG Technologies. http://www.chevron.com/products (accessed Aug 2006). (43) Mukherjee, U. K.; Louie, W. S.; Dahlberg, A. J. Hydrocracking Process for the Production of High Quality Distillates from Heavy Gas Oil. U.S. Patent 6,797,154, 2004.


AlVarez et al.

Figure 3. Schematic representation of the quench zone model.

Figure 4. Effect of feed temperature on the quench requirement, catalyst cycle life, and quench location along the reactor.45

cess;38 (4) atmospheric/vacuum residua upgrading processes according to Chevron’s RDS/VRDS42 and UOP’s RCD Unionfining;38,39 (5) lube oil hydrotreatment (HDT-LUB) such as Exxon’s integrated HYC/HDT37,44 and Chevron’s hydrofinishing42 processes. Details of the mechanical aspects of quench systems as well as fluid temperatures, compositions, and flows within the quench chamber are not available due to the confidential nature of the information. Despite this, some basic information about gas quenching can be found in publications related to hydrotreatment reactor modeling and process simulation. Shah et al.45 analyzed the optimum hydrogen quench location in a residual oil trickle-bed HDS reactor in order to achieve a maximum catalyst cycle life. The reactor model took into account a time-dependent catalyst activity which decreased due to metals deposition. The quench zone was modeled as a stream mixer as shown in Figure 3. The following expressions describe the quench zone mass and energy balances assuming liquid vaporization is insignificant:

q + lout + gout ) lin + gin Tout + TQK Tin ) K+1

(1)

(2)

Figure 5. Effect of quench fluid temperature on the quench requirement, catalyst cycle life, and quench location along the reactor.45

are defined with respect to the feed temperature (TF) in a dimensionless form as:

θM )

TM - TF TF

(4)

θQ )

TQ - TF TF

(5)

Equation 1 expresses the mass balance between the liquid hydrocarbon and gas leaving the previous catalytic bed (lout and gout) and the corresponding streams entering to the following catalytic bed (lin and gin) after mixing with the quench fluid (q). Equation 2 is for determining the inlet temperature to the following catalytic bed (Tin) from the outlet temperature of the previous bed (Tout), quench fluid temperature (TQ), and quenchto-reactant mixture ratio (K). The value of K is calculated from the mass heat capacities (Cp) and mass flow rates of the corresponding streams. In this case, the authors considered that heat capacities do not depend on temperature.

where θM and θQ are the maximum and quench dimensionless temperatures, respectively. The quench requirement is represented by K from eq 3. From Figure 4, it can be seen that there is a strong dependence of the catalyst’s maximum cycle life on the feed temperature and the quench location along the reactor. Increasing the feed temperature (decreasing θM for fixed TM) requires a quench location nearer to the inlet of the reactor in order to maximize the catalyst’s cycle life; if the quench is located after such a position, the cycle life drops substantially. Also, the increase of the mentioned variable increases the quench requirement at any position along the reactor. These results suggest that the feed temperature has to be as low as possible for a required conversion in order to obtain the longest catalyst cycle life. This is explained by the dependency of the catalyst cycle life on the activation energy of the deactivation reaction (i.e., HDM and coking). The high values of activation energy for metals removal

Figures 4 and 5 show the effect of feed and quench temperatures, respectively, as well as quench position along the reactor on the quench requirement and catalyst cycle life. The maximum allowable temperature (TM) and quench temperatures

(44) Carroll, M. B.; Schleicher, G. P.; Boyle, J. P. Integrated Lubricant Upgrading Process. U.S. Patent 6,569,313, 2003. (45) Shah, Y. T.; Mhaskar, R. D.; Paraskos, J. A. Ind. Eng. Chem., Process Des. DeV. 1976, 15, 400-406.

K)

qCQp loutCLp + goutCGp

(3)



Table 1. Comparison of Cycle Lives of Systems with Double and Single Quenches with the Same Total Quench Requirement (K ) 0.15)46 K

z/L

Q1

Q2

0.05 0.05 0.05 0.09 0.09 0.09

0.10 0.10 0.10 0.06 0.06 0.06

0.15

Q1

Q2

catalyst cycle life (d)

Double Quench 0.06 0.485 0.10 0.435 0.20 0.365 0.06 0.875 0.10 0.775 0.20 0.520

203 202 190 205 209 195

Single Quench 0.275

162

give temperature sensitive rate constants; therefore, lower feed temperatures and quench positions closer to the reactor inlet are required in order to reduce the deactivation rate. In the case of Figure 5, although increasing the quench temperature (increasing θQ) also increases the required amount of quench, the catalyst cycle life does not vary significantly. Mhaskar et al.46 extended the previous study into a system with multiple gas quenching. The authors showed that for the same total amount of quench gas, two quench streams can increase the catalyst cycle life compared with a single quench system. A comparison of the cycle lives of systems with double and single quenches with the same total quench requirement (K ) 0.15) is presented in Table 1. It can be seen how the catalyst cycle life increases considerably using two quench streams, especially when increasing the amount of the first quench (Q1) and locating it near the entrance of the reactor. For instance, at values of K ) 0.09 and z/L ) 0.1 for Q1, a cycle life of 209 d is obtained against the 162 d obtained using a single quench. These results are of course of a particular case, and definition of the optimal quench location and number must take into consideration the heat of reactions released as a function of the reactor length and the relative rates of reaction and deactivation. For instance, a feed with a high concentration of aromatics and unsaturated compounds will generate a large amount of heat near the inlet of the reactor, while feeds with lesser amounts of aromatics will release heat more smoothly along the reactor, requiring a lower number of quenches. Yan28 studied the dynamics of a second stage gasoil hydrocracking reactor with and without a quench system. The author recommended the use of transient temperature and product distribution profiles to determine the best quench location, starting time, and amount of quench fluid for proper temperature control, especially during unit startup with fresh catalyst. Using transient profiles allows for observing the reactor temperature response after quenching; in this way, temperature peaks can be avoided by setting a suitable quench position and starting time. For instance, the quench injection starting time during startup must be as soon as or before the hydrocarbon front reaches the quench zone, otherwise temperature peaks will appear leading to hot spot formation. These details cannot be noticed with steady-state profiles, which may result in an inadequate quench system design leading to unsafe operation and temperature runaway during startup. These other two main reasons are well-known to be causes of runaway: (1) wrong design that generates a local shortage or excess of hydrogen (before and after the quench box), for a particular composition of the hydrocarbon in these regions) and (2) lack of balance between locally produced and transfered heat and backmixing (46) Mhaskar, R. D.; Shah, Y. T.; Paraskos, J. A. Ind. Eng. Chem. 1978, 17, 27-33.

in the area. As in the work of Shah et al.,45 the author stressed out the importance of keeping the feed temperatures low, especially for fresh, active catalyst which is more prone to form hot spots. Lababidi et al.47 presented a simulation of an industrial ARDS unit (38 000 bbl/d capacity) with four fixed-bed reactors in series and three interstage hydrogen quench streams. The reactor model took into account HDS, HDM, and coke formation reactions in order to evaluate the effect of catalyst deactivation. The axial temperature, coke, and metal deposition profiles for operation periods of 600 and 2000 h at fixed quench rates are presented in Figure 6a. It can be observed that the temperature rise along the reactor drops as the operation time increases due to catalyst deactivation. Coke and metal deposition on catalyst profiles have discontinuities caused by gas quenching (Q1, Q2, and Q3) between reactors, which implys that the rate of such reactions has a strong dependence on temperature. Van Hasselt et al.34 compared the performance of a traditional cocurrent trickle-bed reactor with a countercurrent reactor configuration in the hydrodesulfurization of a vacuum gasoil. The reaction system included one multibed reactor followed by a single-bed reactor where the number of quench points was adjusted according to the configuration type. For the traditional fixed-bed reactor, four beds with three quench points were employed. The quench zone balances described by eqs 1-3 were extended to take into consideration the mass transfer between the gas and liquid phases. The following expressions represent the quench zone mass balances of H2/H2S in both phases:

Gas:

G G xi,Q q

+

G xi,out gout

-

(

kLi a

(

L Liquid: xi,out lout + kLi a

)

pGi G - CLi MiVmix ) xi,in gin Hi (6)

)

pGi L - CLi MiVmix ) xi,in lin Hi

(7)

Profiles of axial temperature and H2S concentration in the liquid phase of the first reactor are shown in Figure 6b. A sudden H2S concentration drop due to gas quenching is observed. This shows how quench hydrogen, in addition to controlling the reactor temperature, reduces the H2S partial pressure and thus increases the mass transfer to the gas phase, which decreases the H2S concentration in the liquid phase. Palmer and Torrisi48 evaluated and analyzed the requirements for revamping a diesel hydrotreater in order to obtain ultralow sulfur diesel. One of the revamp strategies was to modify the single-bed reactor by dividing the total catalyst volume into two beds and adding a quench zone. Adding a quench zone improved the product quality due to the reduction of the H2S partial pressure; nevertheless, it required reactor volume that otherwise would be occupied by catalyst. Another positive effect of adding gas quench was that the overall heating duty decreased because the quench stream bypassed the charge heater. Bhaskar et al.49 simulated an industrial gasoil HDT unit with two reactors in series and three quenches. Considering the heats of reaction of HDS, HDA, and olefin saturation in the energy balance, the simulated bed effluent temperatures were in excellent agreement with the actual ones as shown in Figure 6c. Unlike in the works of Lababidi et al.47 and Van Hasselt et (47) Lababidi, H. M. S.; Shaban, H. I.; Al-Radwan, S.; Alper, E. Chem. Eng. Technol. 1998, 21, 193-200. (48) Palmer, R. E.; Torrisi, S. Pet. Technol. Q. 2003, ReVamps, 15-18. (49) Bhaskar, M.; Valavarasu, G.; Sairam, B.; Balaraman, K. S.; Balu, K. Ind. Eng. Chem. Res. 2004, 43, 6654-6669.


AlVarez et al.

Figure 6. Hydroprocessing reactor simulations. (a) Axial temperature, COC (coke-on-catalyst) and MOC (metal-on-catalyst): (-) 600; (---) 200 h.47 (b) Axial temperature and H2S concentration in the liquid phase.34 (c) Axial temperature: (-) simulated; (9) actual.49

al.,34 where all quench flowrates were equal, each quench flowrate was adjusted for achieving a bed inlet temperature equal to the feed temperature (340 °C). Stefanidis et al.50 proposed a methodology for estimating a representative operating temperature for simulating an industrial HDS reactor with quench zones. The quench zone model involved an energy balance with a specific heat capacity dependent on temperature for estimating the temperature of the fluid after quench or the quench rate at a fixed exit temperature. The energy balance is as follows:

∫TT

in

out

loutCLp dT +

∫TT

in

out

goutCGp dT +

∫TT

in

Q

qCQp dT ) 0 (8)

Munõz et al.31 studied different alternatives for optimizing the quench flowrate and location during upgrading of heavy oils (50) Stefanidis, G. D.; Bellos, G. D.; Papayannakos, N. G. Fuel Process. Technol. 2005, 86, 1761-1775.

via HDT under moderate reaction conditions. The process was designed with two reactors in series each one with three catalytic beds. With extensive experimental information obtained at the pilot plant scale,51 a reactor model was developed1,52 and scaleup of the process was performed.53 One of the problems suffered when designing the quench system was the availability of recycle hydrogen for quenching the reaction mixture. The procedure included determining the number and position of hydrogen quenches from the recycle stream by keeping the design H2/oil ratio (∼5000 SCF/bbl) reported at experimental scale. To calculate the number of hydrogen quench streams in each reactor, it was necessary to make an overall hydrogen (51) Ancheyta, J.; Betancourt, G.; Marroquıń, G.; Centeno, G.; Alonso, F.; Munõz, J. A. D. Process for the catalytic hydrotreatment of heavy hydrocarbons of petroleum. U.S. Patent pending. (52) Rodrı´guez, M. A.; Ancheyta, J. Energy Fuels 2004, 18, 789-794. (53) Munõz, J. A. D.; Elizalde, I.; Ancheyta, J. Fuel 2007, 86, 12701277.



Figure 7. Quench alternatives of the heavy oil hydrotreatment process. Table 2. H2/Oil Ratio of the Developed Balances31 number and position of hydrogen quenches

H2/oil ratio

case

1st reactor

2nd reactor

1st reactor (ft3 std/bbl)

2nd reactor ft3std/bbl

average (ft3std/bbl)

absolute error (%)

1 2 3 4 5 6 7 exp

2 2 0 1 (1st bed) 1 (2nd bed) 1 (1st bed) 1 (2nd bed) without quench

2 0 2 1 (1st bed) 1 (2nd bed) 1 (2nd bed) 1 (1st bed) without quench

2688 4835 3499 4314 3978 4318 3961 4665

4277 5072 4288 4838 4504 4505 4835 4743

3483 4954 3894 4576 4241 4411 4398 4704

25.9 5.3 17.2 2.7 9.8 6.2 6.5 0

balance, taking into account hydrogen consumption, hydrogen for quench, purges, and hydrogen in other streams. With the determined hydrogen balance, the hydrogen-to-oil ratio for each reactor was calculated and compared with the experimental value. Seven hydrogen balances were determined by changing the number of hydrogen quenches and their position along the reactors. The determined H2/oil ratios together with the experimental value are presented in Table 2. As is seen, balance 4 is the quench hydrogen configuration (Figure 7) that better represents the experimental H2/oil ratio with an absolute error less than 3%. Using more quench streams, for instance balances 1 and 3, decreases substantially the H2/oil ratio; thus, poorer quality of the product would be obtained, and catalyst deactivation due to coke deposition would be increased. The rest of the bed effluents were cooled by heat exchange with other process streams (Figure 7). As an additional exercise, heat integration of the process scheme with and without hydrogen quenches was carried out. It was observed that using hydrogen quench streams reduced the heating utilities and decreased the number of required heat exchangers, which is in agreement with the findings of Palmer and Torrisi.48 2.2. Quenching with Liquids. Different from quench gas based processes, those that use a liquid quench are not so common. That is why most of the information reported in the literature is related to hydroprocessing reactors with hydrogen quench systems. However, quench hydrogen is not always the best option due to its availability in refineries and compression requirements. In such cases, quench liquids may become more attractive due to their higher heat capacity and lower compression costs; nevertheless, it may require more reactor volume or

lower liquid hourly space velocity, i.e., more reactor volume, to achieve the same conversion. The way a liquid quench is introduced into the reactor is different than that of a gas quench (Figure 2), and special reactor internals and liquid quench injection devices are needed to have efficient contact between gas and liquid phases. The processes that use liquid quench streams may be classified into two general categories: (1) multiple feed processes and (2) product recycle processes. In the following sections, a brief description of each type of process will be presented and the effect of the liquid quench on the reaction system will be discussed. 2.2.1. Multiple Feed Processes. Processes with multiple feeds are characterized by introducing several liquid hydrocarbon streams of different composition and properties at the top and between the beds of the reactor. Generally, the hydrocarbon feed is previously fractionated, then the heaviest fraction is fed at the top of the reactor, and lighter fractions are introduced as side feed (Figure 8). By this approach, the side feeds act as quench streams and at the same time are processed together with each bed effluent in the following catalytic bed. Antezana et al.54 proposed a process for hydrotreating distillates with different compositions in a multibed reactor. According to the process scheme, each bed effluent was cooled by direct contact with a second hydrocarbon stream. If the side feed did not provide the desired quench effect, supplementary quench was supplied via heat exchange or by injecting a water (54) Antezana, F. J.; Hochman, J. M.; Weinberg, H. Selective Hydrotreating of Different Hydrocarbonaceous Feedstocks in Temperature Regulated Hydrotreating Zones. U.S. Patent 3,728,249, 1973.


AlVarez et al.

Figure 8. General representation of a multiple feed process.

stream along the side feed as shown in Figure 9. Although water is not desirable in hydroprocessing reactors due to its corroding effect, the authors affirm that in some particular cases it improves the catalyst activity. According to the proposed process scheme,54 the authors presented a comparison of a single-stage conventional HDT (total feed at top) and a multiple feed HDT. In the conventional HDT, the total feed (TF) is a fraction (204-371 °C) containing 8.84 wt % hydrogen; while in the multiple feed scheme, the total feed was previously separated into a heavy fraction (HF) (277-371 °C) and a light fraction (LF) (204-277 °C), and both fractions were used as top feed and side feed. The reactions were carried out under the following operating conditions: T ) 370 °C, P ) 9 MPa, LHSV ) 1 h-1, and H2/oil ratio ) 5000 SCF/bbl. In the multiple feed scheme, the HF and LF were introduced as side feed in the reactor when 50% (case 1) and 25% (case 2) of the total required hydrogen was consumed. Table 3 shows the hydrogen distribution in the liquid feeds (TF, HF, and LF) and in the fractionated products after conventional HDT and multiple feed HDT; additionally, the fractions that were used as the main feed and side feed are specified.54 As can be observed, the product hydrogen content in a fraction used as the main feed is higher than that in the respective fraction coming from the conventional scheme, for instance, the product hydrogen content in the HF of cases 1 and 2 (9.63 and 9.48, respectively) when using the same one as main feed is higher than that in the HF of the conventional scheme (9.37). On the contrary, the product fraction that is used as side feed always has less hydrogen than the respective one of the conventional scheme. This is explained by the contact time of each feed in the catalytic bed; the fraction that is fed at the top has more contact time and thus undergoes more hydrogenation than the fraction that is used as side feed. Since HF and LF have different boiling temperatures, their Cp values are also different, i.e., CpLF > CpHF. This also affects the heat balance and changes the vaporization. These results show how the hydrogen distribution in both fractions varys by feeding different fractions to the reactor and by choosing different locations along the reactor for injecting such feeds. Another example of multiple feed processes is Mobil Oil’s gasoline upgrading process.55 This technology enables to process gasoline, specifically FCC naphtha, by desulfurization in a way that saturation of olefins is reduced. This is accomplished by fractionating the feed and introducing the fractions into the HDT reactor in descendent order of boiling ranges from the top of the reactor to the last bed at the bottom. Since the sulfur containing compounds tend to be more concentrated in the (55) Fletcher, D. L.; Hilbert, T. L.; Sarli, M. S.; Shih, S. S. Gasoline Upgrading Process. U.S. Patent 5,290,427, 1994.

Figure 9. Multiple feed hydrotreatment process according to Antezana et al.: (---) supplemental quench.54 Table 3. Hydrogen Distribution in Liquid Feeds and Product54 product multiple feed feed main Feed side Feed HF, H2 wt % LF, H2 wt % TF, H2 wt %

conventional TF

8.29 9.45 8.84

9.37 10.68

case 1 HF LF 9.63 10.33

LF HF 9.13 10.9

case 2 HF LF 9.48 10.53

LF HF 9.27 10.76

Figure 10. Mobil Oil’s gasoline upgrading process.

heaviest fractions and the olefins in the lightest fractions, the process scheme provides an extended contact time to fractions which require more severe conditions meanwhile the fractions requiring short contact time are introduced toward the end of the reactor reducing olefin saturation. The introduction of olefinpoor fractions at the beginning of the reactor will also limit the temperature rise by reducing the exothermic olefin hydrogenation. Of course, the side feeds are introduced at low temperatures in order to act as quench streams. Since the reaction occurs in the gas phase, the liquid feed is vaporized and contributes to the heat balance based on its heat of vaporization. After HDT, the product is sent to an octane recovery stage. The process scheme is shown in Figure 10. Table 4 shows an example of the operation of the hydrotreatment stage according to the proposed process scheme for two types of FCC naphthas.55 The following reactions conditions were employed: T ) 370 °C, P ) 4.2 MPa, LHSV ) 0.84 h-1, and H2/oil ratio ) 3200 SCF/bbl. It can be observed that substantial sulfur and nitrogen removals are achieved while the octane number is slightly reduced. Unfortunately, there are no experimental data introducing the total feed at the top of the reactor for establishing a comparison; however, it could be expected that in conventional HDT the octane loss would be even greater than in the reported process.



Table 4. Upgrading of FCC Naphtha Fractions: HDT Stage55 naphtha 1 boiling range (°C) sulfur (ppmw) nitrogen (ppmw) research octane motor octane sulfur (ppmw) nitrogen (ppmw) research octane motor octane

Feed 35-260 3800 44 93.5 81.6 Product

Comparison of Quench Systems in Commercial Fixed-Bed

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