Hydrocracking operation with feedstock flexibility

Abstract

A hydrocracking process for converting a petroleum feedstock to lower boiling point products. The process comprises hydrotreating the petroleum feedstock in a pretreatment zone in the presence of hydrogen to produce a hydrotreated effluent stream comprising liquid products. At least a portion of the hydrotreated effluent stream is then passed to a MMS catalyst zone and then to a hydrocracking zone. In one embodiment, the MMS catalyst zone comprises a self-supported multimetallic catalyst prepared from precursors in oxide or hydroxide form. The hydrotreating power in the pretreatment zone is maintained at a level of at least 56%.

Description

[Technical Field]

[0001] This application relates to a reactor system and its use in hydrocarbon reforming. More specifically, this application relates to a two-stage hydrocracking reactor system that extends continuous run time through control of work percentage in the two stages while also allowing for feedstock flexibility. [Background technology]

[0002] A refinery's flexibility and responsiveness to market trends and the regulatory environment significantly impacts its competitiveness. Several factors drive this demand for responsiveness, including the availability of cheap opportunity crude oil and compatible cutter stocks, increasing regulations on residual fuel oils, and price differentials between petrochemical feedstocks, base oils, and transportation fuels. Tighter specifications for refinery process schemes, coupled with more robust catalyst systems, drive a more sustained shift in portfolios with a larger proportion of opportunity feedstocks to a product list that is more aligned with market trends.

[0003] Refineries impose operational constraints to maximize operational reliability. Recent process and catalyst options have been developed that significantly reduce and improve these constraints. As light and heavy crude oil production increases and medium crude oil declines, an increasing number of refineries are supplying opportunity blends of light and heavy crude oils. These crude oil blends raise compatibility concerns, can cause problems in distillation trains, and often exacerbate residual oil contamination in hydrocracker feedstocks. Even when contamination is low enough to approach the detection limits of standard analytical techniques, contaminated residual oil has a detrimental effect on hydrocracker performance. If capital is available, it is possible to invest in improved process options to improve hydrocracker feedstocks, thereby reducing exposure to the adverse effects of opportunity crude oils. Illustrating the current urgency of the need to address compatibility issues, solutions such as distillation and absorption of problematic components are still in practice today, long after they were first proposed. A capital neutral solution is a catalyst system and overall system that can mitigate the risks associated with only a small increase in the final boiling point of the feed to the hydrocracker.

[0004] A hydrocracking operation typically consists of a hydrotreating zone followed by a hydrocracking zone. The goal of the hydrotreating zone is to selectively minimize inhibitors entering the subsequent hydrocracking zone by hydrogenation (or "saturation"). The goal of the hydrocracking zone is to reduce the boiling point of the feedstock, i.e., hydroconvert VGO to lower boiling transportation fuels. If a feedstock other than pure gas-to-liquid wax is used, this further increases the degree of hydrogenation of the material entering the hydrocracking zone.

[0005] The hydrotreating zone reduces hydrocracking inhibitor concentrations by hydrogenating aromatics and organic nitrogen compounds. Because organic nitrogen is relatively easy to quantify, the hydrotreating zone is typically run to an organic nitrogen target. Because simultaneous hydrotreating and hydrocracking produce more refractory compounds, such as dealkylated (nitrogen-containing) aromatics, it is important that hydrocracking activity in the hydrotreating zone be maintained, typically at a minimum of 20-30% VGO conversion.

[0006] The hydrocracking zone reduces the concentration of unconverted oil by lowering the boiling range of the feedstock. The hydrocracking zone is typically run on a VGO hydroconversion target.

[0007] A key issue is how to adjust performance targets (by adjusting process conditions) when the feedstock changes dramatically, for example, when moving from a relatively expensive straight-run VGO to a less expensive pre-treated VGO feedstock. The answer to this question is crucial to a refinery's bottom line. Knowing the proper process conditions has a significant impact on the continuous run time and, therefore, on the refinery's profits. Furthermore, this problem is becoming more prevalent as refineries become exposed to more dynamic crude and cargo markets, and as a result, refineries are now more inclined to accept more dramatic changes in the quality of the feedstock going to the hydrocracker. This dramatically more dynamic crude market requires dramatically more flexible hydrocracker operation in order for refineries to remain competitive. Summary of the Invention

[0008] A hydrocracking process for converting a petroleum feed into lower boiling products in a hydrocracking unit is provided, the process comprising a self-supported mixed metal sulfide (MMS) catalyst. The process includes hydrotreating the petroleum feed in the presence of hydrogen in a pretreatment zone, often having two or more hydrotreating zones, to produce a hydrotreated effluent stream comprising liquid products. At least a portion of the hydrotreated stream effluent from the pretreatment zone is sent to an MMS (mixed metal sulfide) catalyst zone, and the resulting effluent is sent from there to a hydrocracking zone, often comprising two or more reaction zones. The hydrotreating power (e.g., rate of hydrogenation) in the pretreatment zone is maintained at a level of at least 56%. It has been found that controlling the work percentage and maintaining the work percentage at or above 56% significantly improves feedstock flexibility, as well as the continuous run time, stability, and economics of the hydrocracking operation.

[0009] The hydrocracking unit configuration can include a single reactor containing all reaction zones, i.e., pretreatment zone, MMS (mixed metal sulfide) catalyst zone, and hydrocracking zone. In another embodiment, layered loading can be used, where two or more catalysts from different zones are present in the same bed. In another embodiment, the configuration includes two reactors with various reaction zones within each reactor. In one embodiment, when two reactors are used, the first reactor embodies the pretreatment zone, and the second reactor embodies the MMS catalyst zone and the hydrocracking zone.

[0010] Furthermore, it has been discovered that using an unsupported polymetallic catalyst prepared from precursors in oxide or hydroxide form in the upper part of the second reactor in a two-reactor hydrocracking system causes the catalyst to undergo reactions that counteract potential settling. The catalytic reactions at that point in the two-reactor system accomplish this by saturating the major feed components before they are stripped into the incompatible aromatic core. This embodiment further improves the run time and economics of the overall hydrocracking operation. [Brief explanation of the drawings]

[0011] [Figure 1] 1 shows a schematic of a one-stage reactor system with a single reactor and once-through partial conversion.

[0012] [Figure 2] 1 shows a schematic of a single-stage reactor system with multiple reactors and complete or partial conversion with recycle.

[0013] [Figure 3] A two-stage system with complete transformation is shown schematically.

[0014] [Figure 4] FIG. 1 is a schematic diagram of a two-stage reactor system with reversed stages in the reactor (R) and complete conversion.

[0015] [Figure 5] 10 is a graph illustrating the importance of maintaining a pretreatment power ratio above 0.56. DETAILED DESCRIPTION OF THE INVENTION

[0016] By utilizing the process and catalytic reactor system of the present invention, even dramatic changes in feedstock can be easily accommodated. The process of the present invention ensures that the degree of hydrogenation and work performed in the pretreatment hydrotreating zone and in the MMS (mixed metal sulfide) catalyst and hydrocracking zones, respectively, are commensurate with the nature of the feedstock. For example, for a conventional straight-run feed (typically less hydrogen-depleted), the pretreatment zone only needs to hydrogenate 1.2 times as much as the MMS catalyst and hydrocracking zone, whereas for a highly cracked feed (typically more hydrogen-depleted), the pretreatment zone needs to hydrogenate 4 times as much as the MMS catalyst and hydrocracking zone. It has been found that, using the catalyst of the present process, the necessary control can be maintained by controlling the temperature of the catalyst bed in each zone.

[0017] The process of the present invention includes a three-zone reactor system. The first zone is a pretreatment zone. The pretreatment zone generally involves hydrotreating the feedstock. Pretreatment can occur in a single hydrotreating bed or in multiple treatment beds. The number of reaction beds in the pretreatment zone can vary. In one embodiment, the number of hydrotreating reaction beds is greater than one, and in another embodiment, the number of reaction / hydrotreating beds is up to four. This number can be 1, 2, 3, 4, or more, depending on the design.

[0018] The hydrotreating in the pretreatment zone is designed to selectively minimize inhibitors, such as organic N, organic S, and aromatics, that enter the subsequent MMS catalyst and hydrocracking zone. This is accomplished through hydrogenation or saturation. Essentially, the pretreatment zone exists to help protect the subsequent catalysts in the subsequent MMS catalyst and hydrocracking zone that may be sensitive to these inhibitors. It has been discovered by the method of the present invention that using a specific power in the pretreatment zone dramatically increases continuous run times despite large changes in feedstock. This power was discovered in conjunction with a specific catalyst system and has been found to be generally applicable.

[0019] The second zone following the pretreatment zone is a mixed metal sulfide (MMS) catalyst zone. The MMS catalyst can be any mixed metal sulfide catalyst and is generally self-supported. In one embodiment, the MMS catalyst is a self-supported multimetallic catalyst prepared from precursors in the oxide or hydroxide form. In a preferred embodiment, the precursor is in the hydroxide form.

[0020] The third zone is a hydrocracking zone. The hydrocracking zone involves hydrocracking the feedstock to lower its boiling point. The hydrocracking zone can contain a single reactor bed or multiple reactor beds. The hydrocracking zone can have several reactor beds and can include other finishing reactor zones, such as hydrodesulfurization or post-treatment zones. The combined hydrocracking zone can also have two or more reactor beds or multiple reactor beds, depending on the design.

[0021] The hydrocracking system configuration can include a single reactor, two separate reactors, with or without recycle. Product distillation and separation should occur only after the hydrocracking zone. The key aspect is the presence of a pretreatment zone containing one or more hydrotreating reaction zones, an MMS catalyst zone, and a hydrocracking zone containing at least one hydrocracking reaction bed. Additional reaction zones, as well as other finishing reaction zones such as a hydrodesulfurization zone, can also be part of the hydrocracking zone. If present, the hydrodesulfurization zone is often the last reaction bed in the hydrocracking zone.

[0022] The process of the present invention takes advantage of the fact that the hydrocracking, MMS catalyst, and hydrotreating reaction zones increase the saturation of the feedstock through hydrogenation. Hydrogenation consumes hydrogen, increasing the reactor temperature. The refiner controls the temperature of the hydrocracking reactor by adding "quench" hydrogen at critical junctures along the feedstock's path through the reactor. The process of the present invention involves continuously adjusting the average reactor bed temperature to meet the nitrogen target of the pretreatment hydrotreating zone and the conversion target of the MMS catalyst and hydrocracking zones, while also ensuring that the hydrogen consumed in hydrotreating and the hydrogen consumed in hydrocracking are commensurate with the hydrogen deficiency in the feedstock.

[0023] Properly limiting hydrogen consumption has been proven to dramatically increase continuous run times despite large feedstock variations in multiple hydrocracking reactor configurations. Because run terminations are the most costly and frequent event in hydrocracking units, extending continuous run times through proper implementation of hydrogen constraints can add millions of dollars to a refinery's bottom line.

[0024] When using an MMS catalyst, particularly one containing a self-supported multimetallic catalyst prepared from precursors in oxide or hydroxide form, it has been found that controlling hydrogen consumption by maintaining a work percentage in the pretreatment zone of at least 56% provides such benefits. The work percentage is based on controlling the total temperature rise (work) of the beds or reaction zones in the pretreatment zone relative to the total temperature rise of the MMS catalyst and hydrocracking zones. The work percentage of the temperature rise in the pretreatment zone must be at least 56% of the total temperature rise (work) of all reaction beds or reaction zones in the hydrocracking system. The total temperature rise of the beds in the pretreatment zone is divided by the total temperature rise (work) of all reaction beds or reaction zones in the entire hydrocracking system, i.e., the total temperature rise of all reaction beds in the pretreatment zone, MMS catalyst zone, and hydrocracking zone. Surprisingly, it has been found that this work percentage of at least 56% provides significant benefits when using the present MMS catalyst. This power control allows for easy control of the entire process and provides the advantages of extended continuous run times along with feedstock flexibility. The temperature rise or change in each reaction bed or reaction zone can be determined by utilizing the linear temperature increase with depth of each catalyst bed.

[0025] In one embodiment, for a single reactor system with all zones within a single reactor, hydrogen consumption in the hydrotreater or pretreatment zone relative to the MMS catalyst and hydrocracking zone is controlled by controlling the ratio of the total temperature rise (work) in the reaction beds or zones in the pretreatment zone to the total temperature rise (work) in all reaction zones or beds in the hydrocracking system. When the MMS catalyst of the present process is used in the MMS catalyst zone, the work percentage in the pretreatment zone should be at least 56%. This ratio expresses the rate of temperature rise (work) in the reaction beds / zones in the pretreatment zone as a percentage of the total temperature rise of all beds / zones in the hydrocracking system.

[0026] In one embodiment, for a two-reactor system, the hydrogen consumption between the pretreatment zone hydrotreater reactor and the MMS catalyst and hydrocracking zone is controlled to a ratio of about 56 / 44 or greater (56% for the pretreatment zone). As discussed above, it has been discovered that this hydrogen consumption can be adequately controlled by controlling the ratio of the total temperature rise (work) of the reaction beds in the pretreatment zone to the temperature rise (work) in all reaction beds in the hydrocracking system. Significant advantages are realized by controlling the power ratio for the pretreatment zone, which is the ratio of the temperature rise or work exhibited by all beds in the pretreatment zone, i.e., hydrotreater reaction zone, to the temperature rise (work) of all catalyst beds in the reactor system, to at least 56%.

[0027] More specifically, when technical services or operators manage the unit, they manage the bed heat generation (measured as the axial temperature rise across the bed and continuously displayed on the reactor control panel). If the catalyst bed is too active, they add more H2 to quench or add more heat to the bed to increase hydroprocessing activity. There are limits to what the operator can accomplish. They face quench limits (the full-open quench valve limits how much H2 can be introduced) and the fact that they can only heat an individual bed so long before adjacent beds become problematic (overheating). If the bed is too long, it will get hotter unless more quench is available. If the catalyst is deactivated, the temperature cannot be increased much. The use of power facilitated increased plant safety. This was achieved when the operator observed the calculated power and, if necessary, adjusted the reactor temperature to keep it above at least 56%. Temperature settings below 56% reduced hydrogenation activity in all reaction zones within the hydrocracking system. This caused premature deactivation of zones within the hydrocracking system. Temperature settings above 56% allowed residual activity. If left unnoticed, power levels much above 60% and the accompanying high heat release in the hydrotreating zones could lead to the initiation of temperature excursions and ultimately to runaway. Therefore, power levels provide an additional safety measure for process control safety with minimal monitoring, especially during weekends and holidays.

[0028] The idea is to set the power before reaching the MMS catalyst zone so that the pretreatment zone performs the minimum necessary hydrogenation work. This is expressed as the power of the pretreatment zone, defined as the ratio of the temperature rise or work exhibited by the pretreatment zone to the temperature rise or work exhibited by all catalyst beds in the reactor system. It has been discovered that the pretreatment zone performs better than at least 56%, more preferably 58%, and most preferably 60% of the work (calculated from the total heat release in the beds throughout the system). While there is no upper limit to the power, performing all the work in the pretreatment zone represents a major design flaw. If the power is too low (less than 56%), the MMS catalyst bed and subsequent reaction zone will rapidly deactivate, i.e., the beds containing these catalysts will rapidly exhibit lower and lower heat release (axial bed temperature rise).

[0029] A better understanding can be obtained by reviewing Figures 1-5. Figure 1 shows a single-stage reactor system. In the single reactor R shown, there can be four reaction zones, two of which include pretreatment zones. The reactor also includes a MMS (mixed metal sulfide) catalyst zone, preferably with a catalyst prepared from precursors in oxide or hydroxide form, and a hydrocracking zone.

[0030] In Figure 1, petroleum feed FF is fed to reactor R. Hydrotreating can occur in two zones in the pretreatment zone. The resulting hydrotreated effluent is then fed to the MMS catalyst zone containing the present catalyst. The effluent from the MMS catalyst zone is then sent to a hydrocracking zone in reactor R. The bottoms from reactor R can then be sent to a distillation column. Light naphtha, heavy naphtha, kerosene, and diesel can be recovered from the column. The bottoms from the column can be sent to the FCC feedstock and / or a portion can be recycled to reactor R.

[0031] It is important to maintain the pretreatment zone work percentage at a value of at least 56% in the reactor. The work percentage is determined by adding the temperature rise of the two beds in the pretreatment zone and dividing that value by the temperature rise across all beds in the reactor. For example, if the temperature rise in the first bed in the pretreatment zone, or the hydrotreating zone, is 15°C and the temperature rise in the second bed in the pretreatment zone, or the hydrotreating zone, is 10°C, the temperature rise or work in the pretreatment zone is 15°+10°, or 25°. If the temperature rise in the MMS catalyst zone is 8°C and the rise in the hydrocracking zone is 10°C, the calculated work percentage is (15+10) / (15+10+8+10), which is equal to 25 / 43=58%. This is a good result, and maintaining such a work percentage provides the significant benefits noted above.

[0032] FIG. 2 shows a reactor system with multiple reactors R1 and R2. The feedstock is introduced into R1, which may include a pretreatment zone. Pretreatment zone R1 may include, for example, four hydrotreating beds or reaction zones. However, this number can vary. Hydrogenation removes most of the heteroatoms. The hydrotreated effluent is then sent from the bottom of R1 to R2. Reactor R2 includes an MMS catalyst zone, which may use two or more beds or reaction zones, and a hydrocracking zone. The number of reaction beds or reaction zones within each zone can vary. The first or upper reaction zone in R2 is an MMS catalyst zone, which preferably contains a catalyst prepared from a precursor in oxide or hydroxide form. It has been found that in this upper level location of the second reactor, the MMS catalyst can also provide the additional benefit of providing a settling effect.

[0033] The effluent from the MMS catalyst zone is then sent to one or more reactor beds in the hydrocracking zone. Depending on the hydrocracking functionality required in reactor R2, more than one hydrocracking catalyst can be used. The bottom bed of reactor R2 often contains a catalytically effective amount of a desulfurization catalyst for more effective desulfurization, in order to reduce the temperature, inhibit the formation of mercaptans, and convert any recombinant mercaptans to inorganic sulfur. The hydrocracked effluent can then be fed, for example, as bottoms from R2, to a distillation column or fractionator, from which various petroleum products, such as naphtha and diesel, are recovered.

[0034] In Figure 2, reactor R1, if the four beds in R1 (the pretreatment zone) experience temperature increases of 10°C, 15°C, 8°C, and 5°C, respectively, then the work in the pretreatment zone is 38°C (10 + 15 + 8 + 5). In reactor R2, if the work or temperature increase in the MMS catalyst zone is 5°C and the work or temperature increase in the two reactor beds in the hydrocracking zone is 10°C and 17°C, respectively, then the work in the MMS catalyst and hydrocracking zones is 32°C. Therefore, the work percentage in the pretreatment zone is calculated as 38 / (38 + 32), or 54%. This is unacceptable because it has been discovered that when using an MMS catalyst, as in the method and system of the present invention, the power in the pretreatment zone should be at least 56%, more preferably 58% or greater, and most preferably about 60%. In such a case, reactor R1 needs to be made more active for hydrogenation so that the power increases to at least 56%.

[0035] Figures 3 and 4 show two-stage reactor systems. The power in the pretreatment zone also applies to the systems of Figures 3 and 4. However, for the power to apply, the feedstock must flow through successive zones without the introduction of extraneous components, such as recycle to S2 in Figures 3 and 4 or fresh feedstock to R in Figure 4. The power applies to any system in which all reaction zones are in a single vessel, and to systems with zones within vessels without inter-vessel fractionation. For example, in Figure 2, the zones extend from R1 through R2 without the interruption or introduction of extraneous components.

[0036] Figure 5 is a graph of feedstock pretreatment and cracking. The graph shows the relative work—temperature rise—in the pretreatment zone and in the MMS catalyst and hydrocracking zones. In this reaction, for the first 530 days, the unit did not sufficiently pretreat the feedstock, so the power in the pretreatment zone was too low, i.e., below 56%, causing a significant decrease in performance. The graph in the figure shows this degradation due to the ever-decreasing heat release in the MMS catalyst and hydrocracking zones. The decrease in heat release indicates a decrease in performance. Surprisingly, as the power in the pretreatment zone increased above 56%, the system adjusted and thus stabilized. Thus, after 540 days, the heat release in both the pretreatment zone and the MMS catalyst and hydrocracking zones is no longer a function of time, and performance is significantly improved.

[0037] Therefore, the power discussed relates the total temperature rise (work) in the pretreatment zone beds to the total temperature rise (work) in all catalyst beds in the hydrocracking system. This is true for all systems in Figures 1-4. Figure 5 demonstrates the importance of maintaining a power rate in the pretreatment zone of at least 56%.

[0038] It has been found that the method of the present invention, which involves controlling the power of the temperature increase in the pretreatment zone to at least 56%, is applicable only when an MMS catalyst zone containing an unsupported multimetallic catalyst, preferably prepared from precursors in the oxide or hydroxide form, is used in the hydrocracking system. In one embodiment, the catalyst is used in the upper bed of the second reactor, i.e., the upper bed in reactor R2 of Figure 2. In one embodiment, the catalyst is most preferably prepared from precursors in the hydroxide form.

[0039] The preferred catalyst used in the present reactor system is an unsupported multimetallic catalyst (bulk catalyst) prepared by sulfiding a precursor catalyst of the formula:

[0040] A v [(M P )(OH) x (L) n y ] z (M VIB O4), in the formula, A is one monovalent cationic species M P is a promoter metal having an oxidation state of +2 or +4 selected from one or more of Group IIA, Group IIB, Group IVA and Group VIII metals (especially Group VIII, such as Ni), L is an organic oxygen-containing (e.g., maleate) M VIB is a Group VIB metal (e.g., one or more of Mo, W)

[0041] An important aspect of the catalyst precursor (before sulfiding) is that it is in the hydroxide form. Unsupported multimetallic catalyst precursors are in the oxide form, e.g., (Ni) a (Mo) b (W) c O z is not useful.

[0042] In one embodiment, L is selected from carboxylates, carboxylic acids, aldehydes, ketones, the enolate form of an aldehyde, the enolate form of a ketone, and hemiacetals, and combinations thereof.

[0043] In one embodiment, A is NH4 + cations such as ammonium ions, other quaternary ammonium ions, organic phosphonium cations, alkali metal cations, and combinations thereof.

[0044] In one embodiment where both molybdenum and tungsten are used as the Group VIB metals, the atomic ratio of molybdenum to tungsten (Mo:W) is in the range of about 10:1 to 1:10. In another embodiment, the ratio of Mo:W is about 1:1 to 1:5. In one embodiment where molybdenum and tungsten are used as the Group VIB metals, the charge neutral catalyst precursor is represented by the formula A v [(M P )(OH) x (L) n y ] z (Mo t W t’ In yet another embodiment where molybdenum and tungsten are used as the Group VIB metals, chromium can be substituted for some or all of the tungsten, with the ratio of (Cr+W):Mo being in the range of about 10:1 to 1:10. In another embodiment, the ratio of (Cr+W):Mo is 1:1 to 1:5. In one embodiment where molybdenum, tungsten, and chromium are the Group VIB metals, the charge neutral catalyst precursor is of the formula A v [(M P )(OH) x (L) n y ] z (Mo t W t’ Cr ’t” O4).

[0045] In one embodiment, the promoter metal M P is at least one Group VIII metal, M P has an oxidation state of +2 and has the formula Av [(M P )(OH) x (L) n y ] z (M VIB The catalyst precursor of (V-2+2z-x*z+n*y*z)=0.

[0046] In one embodiment, the promoter metal M P is a mixture of two Group VIII metals, such as Ni and Co. In yet another embodiment, M P is a combination of three metals such as Ni, Co and Fe.

[0047] M P In one embodiment, where is a mixture of two Group IIB metals, such as Zn and Cd, the charge neutral catalyst precursor is represented by the formula A v [(Zn a CD a’ )(OH) x (L) y ] z (M VIB O4). In yet another embodiment, M P is a combination of three metals such as Zn, Cd and Hg, and the charge-neutral catalyst precursor is represented by the formula A v [(Zn a CD a’ Hg a” )(OH) x (L) n y ] z (M VIB O4).

[0048] M P In one embodiment, where is a mixture of two Group IVA metals, such as Ge and Sn, the charge neutral catalyst precursor is represented by the formula A v [(Ge b ,Sn b’ )(OH) x (L) n y ] z (M VIB O4). M PIn another embodiment, where is a combination of three Group IVA metals, such as Ge, Sn, and Pb, the charge neutral catalyst precursor has the formula A v [(Ge b Sn b’ Pba b” )(OH) x (L) n y ] z (M VIB O4).

[0049] Promoter metal component M P In one embodiment, the promoter metal (M P The source of the Promoter metal compound is in solution, with the entire amount of the Promoter metal compound dissolved in the liquid to form a homogeneous solution. In another embodiment, the source of the Promoter metal is partially present as a solid and partially dissolved in the liquid. In a third embodiment, the source of the Promoter metal is entirely in solid state.

[0050] Promoter metal compound M P may be a metal salt or mixture of metal salts selected from nitrates, hydrated nitrates, chlorides, hydrated chlorides, sulfates, hydrated sulfates, carbonates, formates, acetates, oxalates, citrates, maleates, fumarates, phosphates, hypophosphites, and mixtures thereof.

[0051] In one embodiment, the promoter metal M P is a nickel compound that is at least partially in a solid state, for example, a water-insoluble nickel compound such as nickel carbonate, nickel hydroxide, nickel phosphate, nickel phosphite, nickel formate, nickel fumarate, nickel sulfide, nickel molybdate, nickel tungstate, nickel oxide, a nickel alloy such as a nickel-molybdenum alloy, Raney nickel, or a mixture thereof.

[0052] In one embodiment, the promoter metal M P is selected from the group IIB and VIA metals such as zinc, cadmium, mercury, germanium, tin or lead and combinations thereof in elemental, compound or ionic form. In yet another embodiment, the promoter metal MP further comprises at least one of Ni, Co, Fe, and combinations thereof in elemental, compound, or ionic form.

[0053] In one embodiment, the promoter metal compound is a zinc compound that is at least partially in a solid state, for example, a zinc compound that is sparingly soluble in water, such as zinc carbonate, zinc hydroxide, zinc phosphate, zinc phosphite, zinc formate, zinc fumarate, zinc sulfide, zinc molybdate, zinc tungstate, zinc oxide, or a zinc alloy such as a zinc-molybdenum alloy.

[0054] In one embodiment, the promoter metal is a Group IIA metal compound selected from the group of water-insoluble compounds such as magnesium compounds, calcium compounds, strontium compounds, and barium compounds that are at least partially in the solid state, e.g., carbonates, hydroxides, fumarates, phosphates, phosphites, sulfides, molybdates, tungstates, oxides, or mixtures thereof.

[0055] In one embodiment, the Promoter metal compound is a tin compound that is at least partially in the solid state, for example, a tin compound that is sparingly soluble in water, such as stannic acid, tin phosphate, tin formate, tin acetate, tin molybdate, tin tungstate, tin oxide, or a tin alloy such as a tin-molybdenum alloy.

[0056] Group VIB metal component: Group VIB metal (M VIBThe Group VIB metal compound may be added in a solid state, a partially dissolved state, or a solution state. In one embodiment, the Group VIB metal compound is selected from molybdenum, chromium, tungsten compounds, and combinations thereof. Examples of such compounds include, but are not limited to, alkali metal, alkaline earth, or ammonium metal oxides of molybdenum, tungsten, or chromium (e.g., ammonium tungstate, ammonium meta-, para-, hexa-, or polytungstate, ammonium chromate, ammonium molybdate, ammonium iso-, peroxo-, di-, tri-, tetra-, hepta-, octa-, or tetradecamolybdate, alkali metal heptamolybdate, alkali metal orthomolybdate, or alkali metal isomolybdate), ammonium salts of phosphomolybdic acid, ammonium salts of phosphotungstic acid, ammonium salts of phosphochromate, molybdenum (di- and tri-)oxides, tungsten (di- and tri-)oxides, chromium or chromic oxide, molybdenum carbide, molybdenum nitride, aluminum molybdate, molybdic acid, chromatic acid, tungstic acid, Mo-P heteropolyanion compounds, Wo-Si heteropolyanion compounds, WP heteropolyanion compounds. The W-Si heteropolyanion compound, Ni-Mo-W heteropolyanion compound, Co-Mo-W heteropolyanion compound or mixtures thereof may be added in a solid, partially dissolved or solute state.

[0057] Chelating Agent (Ligand) L: In one embodiment, the catalyst precursor composition includes at least one non-toxic organic oxygen-containing ligand having an LD50 ratio (as a single oral dose to rats) of greater than 500 mg / Kg. In a second embodiment, the organic oxygen-containing ligand L has an LD50 ratio of greater than 700 mg / Kg. In a third embodiment, the organic oxygen-containing chelating agent has an LD50 ratio of greater than 1000 mg / Kg. As used herein, the term "non-toxic" means that the ligand has an LD50 ratio (as a single oral dose to rats) of greater than 500 mg / Kg. As used herein, the term "at least one organic oxygen-containing ligand" means that in some embodiments, the composition may have two or more organic oxygen-containing ligands, and some of the organic oxygen-containing ligands may have an LD50 ratio of less than 500 mg / Kg, but at least one of the organic oxygen-containing ligands has an LD50 ratio of greater than 500 mg / Kg.

[0058] In one embodiment, the oxygen-containing chelating agent L is selected from the group of non-toxic organic acid addition salts such as formic acid, acetic acid, propionic acid, maleic acid, fumaric acid, succinic acid, tartaric acid, citric acid, oxalic acid, glyoxylic acid, aspartic acid, alkanesulfonic acids such as methanesulfonic acid and ethanesulfonic acid, arylsulfonic acids such as benzenesulfonic acid and p-toluenesulfonic acid, and arylcarboxylic acids such as benzoic acid. In one embodiment, the oxygen-containing chelating agent L is maleic acid (LD of 708 mg / kg).

[0059] In another embodiment, the non-toxic chelating agent L is selected from the group consisting of glycolic acid (having an LD of 1950 mg / kg), lactic acid (LD of 3543 mg / kg), tartaric acid (LD of 7500 mg / kg), malic acid (LD of 1600 mg / kg), citric acid (LD of 5040 mg / kg), gluconic acid (LD of 10380 mg / kg), methoxyacetic acid (LD of 3200 mg / kg), ethoxyacetic acid (LD of 1292 mg / kg), malonic acid (LD of 1310 mg / Kg), succinic acid (LD of 500 mg / kg), fumaric acid (LD of 10700 mg / kg) and glyoxylic acid (LD of 3000 mg / kg). In a further embodiment, the non-toxic chelating agent is selected from the group of organosulfur compounds including, but not limited to, mercaptosuccinic acid (LD50 of 800 mg / kg) and thiodiglycolic acid (LD50 of 500 mg / kg).

[0060] In yet another embodiment, the oxygen-containing ligand L is a carboxylate-containing compound. In one embodiment, the carboxylate compound contains one or more carboxylate functional groups. In yet another embodiment, the carboxylate compound includes monocarboxylates, including but not limited to formate, acetate, propionate, butyrate, pentanoate, and hexanoate, and dicarboxylates, including but not limited to oxalate, malonate, succinate, glutarate, adipate, malate, maleate, fumarate, and combinations thereof. In a fourth embodiment, the carboxylate compound includes maleate.

[0061] The organic oxygen-containing ligand can be mixed with a Promoter metal-containing solution or mixture, a Group VIB metal-containing solution or mixture, or a combination of a Promoter metal and a Group VIB metal-containing precipitate, solution, or mixture. The organic oxygen-containing ligand can be in solution, with the entire amount of the organic oxygen-containing ligand dissolved in a liquid such as water. The organic oxygen-containing ligand can be partially dissolved and partially solid during mixing with the Promoter metal, Group VIB metal, and combinations thereof.

[0062] Diluent Component: The term diluent may be used interchangeably with binder. The use of a diluent is optional in the preparation of the catalyst precursor.

[0063] In one embodiment, a diluent is included in the process for producing the catalyst precursor composition. Typically, the added diluent material has lower catalytic activity than the catalyst prepared from the catalyst precursor composition (without the diluent), or no catalytic activity at all. As a result, in one embodiment, adding a diluent can reduce the activity of the catalyst. Therefore, the amount of diluent added in the process generally depends on the desired activity of the final catalyst composition. Depending on the intended catalyst application, a diluent amount of 0 to 95 wt. % of the total composition may be appropriate.

[0064] The diluent can be added simultaneously or alternately to the promoter metal component, promoter metal-containing mixture, Group VIB metal, or metal-containing mixture. Alternatively, the promoter metal and Group VIB metal mixture can be combined together, and then the diluent can be added to the combined metal mixture. It is also possible to combine a portion of the metal mixture simultaneously or alternately, then add the diluent, and finally add the remainder of the metal mixture simultaneously or alternately. Furthermore, it is also possible to combine the diluent with the metal mixture in a solute state, and then add at least a partial metal compound in a solid state. An organic oxygen-containing ligand is present in at least one of the metal-containing mixtures.

[0065] In one embodiment, the diluent is combined with the Group VIB metal and / or promoter metal prior to being combined with the bulk catalyst precursor composition and / or prior to being added during preparation of the bulk catalyst precursor composition. In one embodiment, combining the diluent with any of these metals is accomplished by impregnation of the solid diluent with these materials.

[0066] Diluent materials include any material conventionally applied as a diluent or binder in hydroprocessing catalyst precursors. Examples include silica, silica-alumina, such as conventional silica-alumina, silica-coated alumina and alumina-coated silica, alumina, such as (pseudo)boehmite or gibbsite, titania, zirconia, cationic or anionic clays, such as saponite, bentonite, kaolin, sepiolite or hydrotalcite, or mixtures thereof. In one embodiment, the binder material is selected from silica, aluminum-doped colloidal silica, silica-alumina, alumina, titanium, zirconia, or mixtures thereof.

[0067] These diluents can be applied as they are or after peptization. It is also possible to apply precursors of these diluents that are converted into any of the above-mentioned diluents during the process. Suitable precursors are, for example, alkali metal or ammonium aluminates (to obtain alumina diluents), water glass or ammonium or acid-stabilized silica sols (to obtain silica diluents), mixtures of aluminates and silicates (to obtain silica-alumina diluents), mixtures of divalent, trivalent and / or tetravalent metal sources, such as mixtures of water-soluble salts of magnesium, aluminum and / or silicon (to prepare cationic clays and / or anionic clays), chlorohydrol, aluminum sulfate, or mixtures thereof.

[0068] Other Optional Components: In addition to the components listed above, other materials, including other metals, can be added if desired. These materials include any materials added during the preparation of conventional hydroprocessing precursors. Suitable examples include phosphorus compounds, boron compounds, additional transition metals, rare earth metals, fillers, or mixtures thereof. Suitable phosphorus compounds include ammonium phosphate, phosphoric acid, or organic phosphorus compounds. The phosphorus compounds can be added at any stage of the process. Suitable additional transition metals that can be added during the process include, for example, rhenium, ruthenium, rhodium, iridium, chromium, vanadium, iron, cobalt, nickel, zinc, platinum, palladium, cobalt, and the like. In one embodiment, the additional metals are applied in the form of water-insoluble compounds. In another embodiment, the additional metals are added in the form of water-soluble compounds. Apart from adding these metals during the process, the final catalyst precursor composition can also be mixed with the optional materials. For example, the final catalyst precursor composition can be impregnated with an impregnation solution containing any of these additional materials.

[0069] Methods for producing hydroprocessing catalyst precursors: Preparation methods allow for systematic variation of the composition and structure of the catalyst precursor by controlling the relative amounts of elements, the types of reagents, and the length and severity of the various reactions and reaction steps.

[0070] The order of addition of the reagents used in forming the catalyst precursor is not critical. For example, the organic oxygen-containing ligand can be combined with a mixture of the Promoter metal and Group VIB metal(s) prior to precipitation or cogelation. The organic oxygen-containing ligand can be mixed with a solution of the Promoter metal and then added to a solution of one or more Group VIB metals. The organic oxygen-containing ligand can be mixed with a solution of one or more Group VIB metals and then added to a solution of one or more Promoter metals.

[0071] Forming a Precipitate or Co-Gel with a Group VIB / Promoter Metal: In one process embodiment, the first step is a precipitation or co-gelation step, which involves reacting a Promoter metal component in solution with a Group VIB metal component in solution in a mixture to obtain a precipitate or co-gel. The precipitation or co-gelation is carried out at a temperature and pH at which the Promoter metal compound and the Group VIB metal compound precipitate or form a co-gel. An organic oxygen-containing ligand, in solution or at least partially in solution, is then combined with the precipitate or co-gel to form one embodiment of a catalyst precursor.

[0072] In one embodiment, the temperature at which the catalyst precursor is formed is between 50 and 150°C. If this temperature is below the boiling point of the protic liquid, such as 100°C in the case of water, the process is generally carried out at atmospheric pressure. Above this temperature, the reaction is generally carried out at increased pressure, such as in an autoclave. In one embodiment, the catalyst precursor is formed at a pressure between 0 and 3000 psig. In a second embodiment, it is between 100 psig and 1000 psig.

[0073] The pH of the mixture can be varied to increase or decrease the rate of precipitation or co-gelation, depending on the desired properties of the product. In one embodiment, the mixture is maintained at its natural pH during the reaction step. In another embodiment, the pH is maintained in the range of 0 to 12. In another embodiment, it is between 4 and 10. In a further embodiment, the pH is in the range of 7 to 10. Changing the pH can be achieved by adding a base or acid to the reaction mixture, or by adding hydroxide ions or H ions, which increase or decrease the pH, respectively, upon increasing the temperature. - This can be done by adding compounds that break down into ions. Examples include urea, nitrite, ammonium hydroxide, inorganic acids, organic acids, inorganic bases, and organic bases.

[0074] In one embodiment, the reaction of the Promoter metal component is carried out with a water-soluble metal salt, such as zinc, molybdenum, and tungsten metal salts. The solution may further contain other Promoter metal components, such as Group VIII metal components including cadmium or mercury compounds such as Cd(NO) or (CHCO)Cd, cobalt or iron compounds such as Co(NO) or (CHCO)Co, and other Group VIB metal components such as chromium.

[0075] In one embodiment, the reaction of the Promoter metal components is carried out with water-soluble tin, molybdenum, and tungsten metal salts. The solution may further contain other Group IVA metal components, for example, lead compounds such as Pb(NO3)4 or (CH3CO2)2Pb, and other Group VIB metal compounds, such as chromium compounds.

[0076] The reaction is carried out with appropriate metal salts to result in the precipitation or cogelation of zinc / molybdenum / tungsten, tin / molybdenum / tungsten, zinc / molybdenum, zinc / tungsten, tin / molybdenum, tin / tungsten, or zinc / tin / molybdenum / tungsten, or nickel / molybdenum / tungsten, cobalt / molybdenum / tungsten, nickel / molybdenum, nickel / tungsten, cobalt / molybdenum, cobalt / tungsten, or nickel / cobalt / molybdenum / tungsten. The organic oxygen-containing ligand can be added before or after the precipitation or cogelation of the promoter metal compound and / or Group VIB metal compound.

[0077] The metal precursor can be added to the reaction mixture in solution, suspension, or a combination thereof. If a soluble salt is added neat, it dissolves in the reaction mixture and then precipitates or co-gels. The solution can be heated, optionally under vacuum, to effect precipitation and evaporation of water.

[0078] After precipitation or cogelation, the catalyst precursor can be dried to remove water. Drying can be carried out under atmospheric conditions or under an inert atmosphere such as nitrogen, argon, or vacuum. Drying can be carried out at a temperature sufficient to remove water but not organic compounds. Preferably, drying is carried out at about 120°C until a constant weight of catalyst precursor is reached.

[0079] Forming a Precipitate with an Optional Binder Component: In one embodiment using a binder, the binder component can be added to a reaction mixture containing the metal precursors in solution, suspension, or a combination thereof to form a precipitate or co-gelation. The precipitate is then dried to remove water.

[0080] In one embodiment using magnesium aluminosilicate clay as a binder, a first reaction mixture is formed that includes a silicon component, an aluminum component, a magnesium component, a promoter metal compound, and / or a Group VIB metal compound. In one embodiment, the first reaction mixture is formed under ambient pressure and temperature conditions. In one embodiment, the reaction is carried out at a pressure ranging from 0.9 bar to 1.2 bar and a temperature ranging from about 0°C to 100°C.

[0081] Examples of silicon components include, but are not limited to, sodium silicate, potassium silicate, silica gel, silica sol, silica gel, hydronium- or ammonium-stabilized silica sol, and combinations thereof. Examples of aluminum components useful in the method of the present invention include, but are not limited to, sodium aluminate, potassium aluminate, aluminum sulfate, aluminum nitrate, and combinations thereof. Examples of magnesium components useful in the method of the present invention include, but are not limited to, magnesium metal, magnesium hydroxide, magnesium halides, magnesium sulfate, and magnesium nitrate. In one embodiment, a sufficient amount of acid is added to a mixture containing a metal precursor and a binder component to adjust the pH of the mixture to about 1 to about 6, forming a first reaction mixture.

[0082] After forming the first reaction mixture, an alkali base is added to form a second reaction mixture. Examples of alkali bases include, but are not limited to, ammonium hydroxide, sodium hydroxide, and potassium hydroxide. Sufficient alkali base is added to the first reaction mixture so that the resulting second reaction mixture has a pH of about 7 to about 12. The second reaction mixture is then reacted for a sufficient time and at a sufficient temperature to form a catalyst precursor incorporating at least the clay as a binder. In some embodiments, the time is at least 1 second. In a second embodiment, it is 15 minutes. In a third embodiment, it is at least 30 minutes. The temperature of the second reaction mixture can range from about 0°C to about 100°C. The reaction can be carried out at ambient pressure, although higher or lower pressures are not excluded.

[0083] In one embodiment using magnesium aluminosilicate clay as the binder, the ratio of silicon to aluminum to magnesium can be expressed as the elemental molar ratio: aSi:bAl:cMg, where "a" has a value of 3-8, "b" has a value of 0.6-1.6, and "c" has a value of 3-6.

[0084] Catalyst Precursor Characterization: Characterization of the charge-neutral catalyst precursor can be performed using techniques known in the art, including, but not limited to, powder X-ray diffraction (PXRD), elemental analysis, surface area measurements, average pore size distribution, and average pore volume. Porosity and surface area measurements can be performed using BJH analysis under BET nitrogen adsorption conditions.

[0085] Properties of the catalyst precursor: In one embodiment, the catalyst precursor has an average pore volume of 0.05 to 5 ml / g as determined by nitrogen adsorption. In another embodiment, the average pore volume is 0.1 to 4 ml / g. In a third embodiment, it is 0.1 to 3 ml / g.

[0086] In one embodiment, the catalyst precursor is at least 10 m 2 In a second embodiment, the surface area is at least 50 m 2 In a third embodiment, the surface area is at least 150 m2 / g of surface area.

[0087] In one embodiment, the catalyst precursor has an average pore size, as determined by nitrogen adsorption, of 2 to 50 nanometers, in a second embodiment, of 3 to 30 nanometers, and in a third embodiment, of 4 to 15 nanometers.

[0088] In one embodiment that includes magnesium aluminosilicate clay as a binder, the catalyst precursor is a layered material composed of stacks of elementary clay platelets.

[0089] Forming Process: In one embodiment, the catalyst precursor composition can be directly formed into various shapes, generally depending on the intended commercial application. These shapes can be made by any suitable technique, such as extrusion, pelletizing, beading, or spray drying. If the bulk catalyst precursor composition contains too much liquid to be subjected to a direct forming process, solid-liquid separation can be performed before forming.

[0090] Addition of Pore-Forming Agents The catalyst precursor can be mixed with a pore-forming agent, including, but not limited to, stearic acid, polyethylene glycol polymers, carbohydrate polymers, methacrylates, and cellulose polymers. For example, the dried catalyst precursor can be mixed with a cellulose-containing material, such as methylcellulose, hydroxypropylcellulose, or other cellulose ethers, in a ratio of 100:1 to 10:1 (wt. % catalyst precursor to wt. % cellulose), and water can be added until an extrudable mixture is obtained. Examples of commercially available cellulose-based pore-forming agents include, but are not limited to, METHOCEL™ (available from DuPont), Avicel® (available from DuPont), and porocel (available from Porocel). The extrudable mixture can be extruded and then optionally dried. In one embodiment, drying can be performed under an inert atmosphere, such as nitrogen, argon, or vacuum. In another embodiment, drying can be performed at an elevated temperature between 70°C and 200°C. In yet another embodiment, drying is performed at 120°C.

[0091] Sulfiding Agent Component: The charge-neutral catalyst precursor can be sulfided to form an active catalyst. In one embodiment, the sulfiding agent is elemental sulfur alone. In another embodiment, the sulfiding agent is a sulfur-containing compound that can decompose into hydrogen sulfide under prevailing conditions. In yet a third embodiment, the sulfiding agent is HS alone or in H.

[0092] In one embodiment, the sulfiding agent is ammonium sulfide, ammonium polysulfide ([(NH4)2S x), ammonium thiosulfate ((NH4)2S2O3), sodium thiosulfate (Na2S2O3), thiourea CSN2H4, carbon disulfide, dimethyl disulfide (DMDS), dimethyl sulfide (DMS), dibutyl polysulfide (DBPS), mercaptans, tertiary butyl polysulfide (PSTB), tertiary nonyl polysulfide (PSTN), and the like. In another embodiment, the sulfiding agent is selected from alkali and / or alkaline earth metal sulfides, alkali and / or alkaline earth metal hydrosulfides, and mixtures thereof. The use of sulfiding agents containing alkali and / or alkaline earth metals may require an additional separation process step to remove the alkali and / or alkaline earth metals from the spent catalyst.

[0093] In one embodiment, the sulfiding agent is ammonium sulfide in an aqueous solution, which can be synthesized from hydrogen sulfide and ammonia refinery off-gas. This synthesized ammonium sulfide is readily soluble in water and can be easily stored in an aqueous solution in a tank prior to use. In one embodiment, sulfiding is also carried out using the aqueous ammonium sulfide solution in the presence of at least one sulfur additive selected from the group consisting of thiodazole, thioacid, thioamide, thiocyanate, thioester, thiophenol, thiosemicarbazide, thiourea, mercaptoalcohol, and mixtures thereof.

[0094] In one embodiment, a hydrocarbon feedstock is used as a sulfur source to sulfiding the catalyst precursor, which can be carried out in one or more hydrotreating reactors during hydroprocessing.

[0095] In one embodiment, the sulfiding agent is present in an amount greater than the stoichiometric amount required to form the sulfided catalyst from the catalyst precursor. In another embodiment, the amount of sulfiding agent corresponds to a sulfur to Group VIB metal molar ratio of at least 3 to 1 to produce the sulfided catalyst from the catalyst precursor. In a third embodiment, the total amount of sulfur-containing compounds is generally selected to correspond to any of about 50-300%, 70-200%, and 80-150% of the stoichiometric amount of sulfur required to convert the metal to, for example, CO9S8, MoS2, WS2, Ni3S2, etc.

[0096] Sulfiding Step: Sulfiding of the catalyst precursor to form the catalyst (sometimes referred to as "presulfiding") can occur prior to introducing the catalyst into the hydroprocessing reactor (hence, ex-situ sulfiding). In another embodiment, the sulfiding is in situ. In one embodiment, the sulfiding process is performed ex-situ, preventing the formation of undesirable compounds within the hydroprocessing unit. In one embodiment, the catalyst precursor is converted to the active catalyst upon contact with a sulfiding agent under H2-containing gas pressure at a temperature ranging from 70°C to 500°C for 10 minutes to 15 days. When the sulfiding temperature is below the boiling point of the sulfiding agent, such as 60-70°C for ammonium sulfide solution, the process is generally carried out at atmospheric pressure. At temperatures above the boiling points of the sulfiding agent / optional components, the reaction is generally carried out at elevated pressure.

[0097] In one embodiment, sulfiding can be carried out in the vapor phase using hydrogen and a sulfur-containing compound decomposable to HS. Examples include mercaptans, CS, thiophenes, DMS, DMDS, and suitable sulfur-containing refinery effluent gases. The use of HS alone is sufficient. Contact between the catalyst precursor in the vapor phase and the hydrogen and sulfur-containing compound can be carried out in one step at temperatures ranging from 125°C to 450°C (257°F to 842°F) in one embodiment, and from 225°C to 400°C (437°F to 752°F) in another embodiment. In one embodiment, sulfiding is carried out by increasing the temperature, for example, at increments of 0.5 to 4°C (0.9 to 7.2°F) per minute, for a period of time, and holding until completion, for example, for 1 to 12 hours.

[0098] As used herein, completion of the sulfurization process means that at least 95% of the stoichiometric amount of sulfur required to convert the metal to, for example, CO9S8, MoS2, WS2, Ni3S2, etc. has been used up.

[0099] In another embodiment of sulfurization in the gas phase, sulfurization is carried out in two or more steps, with the first step being at a lower temperature than the subsequent steps. For example, the first step is carried out at about 100-250°C (212°F-482°F), preferably about 125-225°C (257°F-437°F). After a short time, e.g., 1 / 2 to 2 hours (the temperature is held constant), the second step can be carried out at about 225-450°C (437°F-842°F), preferably about 250-400°C (482°F-752°F). The total pressure during the sulfurization step can be from atmospheric pressure to about 10 bar (1 MPa). The gaseous mixture of H2 and sulfur-containing compound can be the same or different in these steps. Gas-phase sulfurization can be carried out by any suitable method, including fixed-bed processes and moving-bed processes (where the catalyst moves relative to the reactor, e.g., boiling processes and rotary kilns).

[0100] In one embodiment, sulfiding is carried out in the liquid phase. The catalyst precursor is first contacted with an organic liquid in an amount ranging from 20 to 500% of the pore volume of the catalyst precursor. Contact with the organic liquid can be at a temperature ranging from ambient temperature to 250°C (482°F). After incorporation of the organic liquid, the catalyst precursor is contacted with hydrogen and a sulfur-containing compound.

[0101] In one embodiment, the organic liquid has a boiling point range of about 100-550°C (212-1022°F). In another embodiment, the organic liquid is a petroleum fraction such as heavy oil, a lubricating oil fraction such as mineral lubricating oil, atmospheric gas oil, vacuum gas oil, straight run gas oil, white spirit, middle distillates such as diesel, jet fuel and kerosene, naphtha, and gasoline. In one embodiment, the organic liquid contains less than 10% by weight sulfur, preferably less than 5% by weight sulfur.

[0102] In one embodiment, sulfiding in the liquid phase (or "start-up") is carried out as a "quick" process, with sulfiding occurring over a period of less than 72 hours and with a temperature increase in the range of 0.5-4°C (0.9-7.2°F) per minute. In a second embodiment, the quick start takes less than 48 hours. In a third embodiment, it takes less than 24 hours.

[0103] In rapid sulfiding, contact between the catalyst precursor in the organic liquid and hydrogen and the sulfur-containing compound can be carried out in one step, in one embodiment at a temperature of 150-450°C, and in another embodiment at a temperature of 225-400°C. In yet another embodiment of rapid sulfiding, sulfiding is carried out in two or more steps, with the first step being at a lower temperature than subsequent steps. For example, the first step is at about 100-250°C (212-482°F), or about 125-225°C (257-437°F). After a short period of time, e.g., 1 / 2 to 2 hours (where the temperature is held constant), the temperature is then increased for the second step, e.g., to 250-450°C (482-842°F), preferably 225-400°C (437-7520°F). The temperature is maintained for 1 to 36 hours, after which the sulfiding is complete.

[0104] In yet another embodiment, liquid-phase sulfiding is carried out as a "slow" process, where sulfiding is carried out over a period of 4 days to up to 3 weeks, i.e., for at least 96 hours. In this slow process, contact between the catalyst precursor in the organic liquid and the hydrogen and sulfur-containing compound is carried out in two or more steps, with the first step being at a lower temperature than subsequent steps, and the temperature being increased slowly, e.g., in hourly increments, rather than minute-by-minute as in the rapid start. The gaseous mixture of H2 and sulfur-containing compound can be the same or different in these steps. In one embodiment, the first step is at about 100-375°C (212°F-707°F), preferably about 125-350°C (257°F-662°F), with a temperature ramp rate of 0.25-4°C (0.45-7.2°F) per hour. After the first step, the temperature is held constant for a period of 2 to 24 hours and then increased at a rate of 5 to 20°C (9 to 36°F) per hour for the second step. In one embodiment, the second step is carried out at about 200 to 450°C (392 to 842°F), preferably about 225 to 400°C (437 to 752°F).

[0105] In one embodiment, sulfiding is carried out using elemental sulfur, which is incorporated into the pores of the catalyst precursor. In this process, elemental sulfur is mixed with the catalyst precursor in an amount of 2-15% by weight of the catalyst precursor at a temperature below the melting point of sulfur. In one embodiment, the mixing is at 180°F to 210°F (82°C to 99°C). Sequentially or simultaneously with mixing of the precursor and elemental sulfur, the mixture is contacted with a high-boiling organic liquid. The mixture is then heated to a temperature in the range of 250-390°F (121°C to 199°C) in the presence of nitrogen to produce HS and metal sulfides. In one embodiment, the organic liquid is selected from the group consisting of olefins, gasoline, white spirit, diesel, gas oil, mineral lubricating oil, and white oil.

[0106] In one embodiment, catalysts sulfided from catalyst precursor embodiments are surprisingly found to provide roughly the same 700°F+ conversion whether sulfided via the gas phase or the liquid phase as in the "fast" process. In one embodiment, the use of catalysts sulfided via the liquid phase and the "slow" process has been found to increase the 700°F+ conversion by at least 25%. In yet another embodiment, the 700°F+ conversion is doubled using catalysts sulfided via the slow process.

[0107] A preferred catalyst precursor is a Ni-Mo-W maleate catalyst precursor. The catalyst is preferably sulfided with dimethyl sulfide (DMDS).

[0108] <Feed material> A wide range of petroleum and chemical feedstocks can be hydroprocessed in accordance with the present invention. Suitable feedstocks include whole crude oil and reduced petroleum residua, atmospheric and vacuum residua, propane deasphalted residua, e.g., bright stock, cycle oil, FCC tower residua, gas oils including atmospheric and vacuum gas oils and coker gas oils, light to heavy distillates including unrefined virgin distillates, hydrocracked products, hydrotreated oils, dewaxed oils, slack wax, Fischer-Tropsch wax, raffinates, naphtha, and mixtures of these materials. Typical lighter feedstocks generally include distillate fractions boiling between about 175°C (about 350°F) and about 375°C (about 750°F). Using such feedstocks, significant amounts of hydrocracked naphtha, which can be used as a low-sulfur gasoline blendstock, are produced. Typical heavier feedstocks include, for example, vacuum gas oils boiling up to about 593°C (about 1100°F), usually in the range of about 350°C to about 500°C (about 660°F to about 935°F), in which case a correspondingly greater proportion of diesel fuel is produced.

[0109] In one embodiment, this process operates by conducting a feedstock containing high levels of sulfur and nitrogen through an initial treatment reaction stage to convert a significant amount of the sulfur and nitrogen in the feedstock to inorganic forms; the primary objective of this step is to reduce the nitrogen content of the feedstock. The hydrotreating step is carried out in one or more reaction zones (catalyst beds) in the presence of hydrogen and a hydrotreating catalyst. The conditions used are suitable for hydrodesulfurization and / or denitrification, depending on the characteristics of the feedstock. The product stream is then sent directly (without separation) or, with separation and water washing, to a hydrocracking zone where boiling range conversion takes place. After the initial hydroconversion step, a bed of bulk multimetallic catalyst may be applied, typically at the bottom of a second reactor (R2). In a two-stage unit, the liquid hydrocarbon stream from the first hydroconversion stage, along with hydrotreat gas and other hydrotreating / hydrocracking reaction products, including hydrogen sulfide and ammonia, travels to a separator in which hydrogen, light ends, inorganic nitrogen, and hydrogen sulfide are removed from the hydrocracked liquid product stream. The recycle gas may be scrubbed to remove ammonia and subjected to an amine wash to remove hydrogen sulfide to improve the purity of the recycled hydrogen and thereby reduce the sulfur level of the product. In the second stage hydroconversion, the hydrocracking reaction is completed. Immediately following the second hydroconversion stage, a bed of bulk multimetallic catalyst may be provided, typically in the bottom of the third reactor (R3). The bulk multimetallic catalyst bed is effective under the R2 and / or R3 reaction conditions for the removal of "hard sulfur" species, i.e., sulfur species having atmospheric boiling points in the range of about 93°C to about 593°C (200°F to about 1100°F), particularly in the range of about 350°C to about 500°C (about 660°F to about 935°F).

[0110] <Hydrotreating Catalysts> Whether in the first or second reactor, conventional hydrotreating (hydrodesulfurization) catalysts can be used in the hydrodesulfurization zone. Typical conventional hydrodesulfurization catalysts for use in the present invention include those composed of at least one Group VIII metal, preferably Fe, Co, or Ni, more preferably Co and / or Ni, most preferably Co; and at least one Group VIB metal, preferably Mo or W, more preferably Mo, on a relatively high surface area support material, preferably alumina. Other suitable hydrodesulfurization catalyst supports include zeolites, amorphous silica-alumina, and titania-alumina noble metal catalysts. Noble metal catalysts can also be used, preferably where the noble metal is selected from Pd and Pt. More than one hydrodesulfurization catalyst can be used in different beds within the same reactor. The Group VIII metal is typically present in an amount ranging from about 2 to about 20 wt. %, preferably from about 4 to about 12 wt. %. The Group VIB metal is typically present in an amount ranging from about 5 to about 50 weight percent, preferably from about 10 to about 40 weight percent, and more preferably from about 20 to about 30 weight percent, with all metal weight percentages being relative to the support (percentages based on the weight of the support).

[0111] <Hydrocracking Catalysts> Examples of additional hydrocracking catalysts that can be used in the hydrocracking stage zone include nickel, nickel-cobalt-molybdenum, cobalt-molybdenum, and nickel-tungsten and / or nickel-molybdenum, with the latter two being preferred. Non-limiting examples of noble metal catalysts include those based on platinum and / or palladium. Porous support materials that can be used for both noble and non-noble metal catalysts include refractory oxide materials such as alumina, silica, alumina-silica, kieselguhr, diatomaceous earth, magnesia, or zirconia, with alumina, silica, and alumina-silica being preferred and most common. Zeolite supports can also be used, particularly large-pore faujasites such as USY.

[0112] Numerous hydrocracking catalysts are available from different commercial suppliers and can be used according to the feedstock and product requirements, and the functionality of the hydrocracking catalyst can be determined empirically. The choice of hydrocracking catalyst is not critical. Any catalyst having the desired hydroconversion functionality at the selected operating conditions can be used, including conventional hydrocracking catalysts.

[0113] The precipitation of molecules that are no longer compatible with gas oils as they are hydroprocessed is a particularly serious problem in hydrocrackers operated with a two-stage recycle. The recycle operation exacerbates the impact of incompatibility because it concentrates polycyclic aromatic solutes by hydrocracking the solvent into transportation fuels and by distilling off these fuels as products. This concentrates aromatic solutes that are already poorly compatible, and the concentration further increases the risk of aromatic solutes exceeding their solubility limits. The risk of catastrophic shutdown due to precipitation in the recycle loop is well known. Precipitation is typically controlled by carefully monitoring the recycle loop for the accumulation of sediment-forming heavy aromatic compounds and by releasing an appropriate proportion of unconverted oil to keep these aromatic compounds below their solubility limits. Several monitoring options exist, including: i) by color and appearance; ii) by UV-Vis spectroscopy, from which a Polycyclic Aromatics Index (or PCI, a marker of incompatibility risk) can be derived; and iii) by spot-checking aromatics accumulation using high-resolution mass spectrometry. This reactive approach to mitigating the accumulation of incompatible aromatics results in a bleed stream of low-value, unconverted oil, resulting in a missed opportunity to hydroprocess this oil into valuable transportation fuels. A more proactive approach to mitigating the risk of incompatible compounds before they begin to accumulate has been demonstrated by using a catalyst of the present invention prepared from a hydroxide form of precursor. This proactive approach minimizes bleed streams and, accordingly, improves transportation fuel production rates without significantly impacting continuous operation time. This proactive approach utilizes a catalyst of the present invention prepared from a hydroxide form of precursor in the first hydrocracking zone of a two-stage system containing two reactors.

[0114] By selectively maintaining polycyclic aromatic compounds in the hydroprocessing stream at their hydrogenation equilibrium before they have a chance to aggregate, emulsify, or otherwise become a nuisance to the hydroprocessing process, precipitation can be proactively counteracted. Because incompatible large aromatic compounds already aggregate at low concentrations, and because incompatible large aromatic compounds boil toward and beyond the end of the boiling range of a typical hydrocracker feedstock, conventional hydroprocessing catalysts focus on saturating the majority of the feedstock (solvent) and do not begin to saturate low concentrations of polycyclic aromatic compounds (solutes) until they have accumulated to significant (and undesirable) levels. In stark contrast, the unique pore structure and high hydrogenation activity of the catalysts of the present invention enable them to selectively target polycyclic aromatic compounds, even at concentrations far below the level at which they begin to aggregate. A comparison of the effectiveness of a catalyst system with the catalyst of the present invention in the first bed position of a second-stage hydrocracker compared to a typical catalyst system demonstrated the selective hydrogenation of representative polycyclic aromatic compounds doped in a typical Middle Eastern vacuum gas oil. The significant reduction in PCI (a marker of incompatibility risk) by adding the catalyst of the present invention to a catalyst system deployed for hydroprocessing of a vacuum gas oil (VGO) feedstock obtained almost exclusively from atmospheric residue desulfurization (ARDS) in a once-through operation also demonstrated the benefit of using the catalyst of the present invention in further destroying sediments.

[0115] The process and reactor system of the present invention provide extended continuous run times, thereby significantly improving the economics of two-stage hydrocracking facilities. By controlling the power, extended continuous run times are achieved while also allowing for great feedstock flexibility. The use of the catalysts of the present invention, particularly those prepared from hydroxide-form precursors, also greatly amplifies these benefits. Furthermore, when the catalysts of the present invention are used in the top level of the second reactor of a two-stage system, sediment destruction is also achieved, thereby further improving continuous run times. The process and reactor system of the present invention result in dramatically more flexible and improved hydrocracker operation. [Example]

[0116] The preparation of catalysts for use in the process of the present invention is illustrated in the following examples, which are intended to be non-limiting.

[0117] [Example 1] Ni-Mo-W-maleate catalyst precursor Formula (NH4){[Ni 2.6 (OH) 2.08 (C4H2O4 2- ) 0.06 ](Mo 0.35 W 0.65 A catalyst precursor for (NH4)MoO was prepared as follows: 52.96 g of ammonium heptamolybdate (NH4)MoO 24Ni(NO3)2·4H2O was dissolved in 2.4 L of deionized water at room temperature. The pH of the resulting solution was in the range of 5-6. 73.98 g of ammonium metatungstate powder was then added to the above solution and stirred at room temperature until completely dissolved. 90 ml of concentrated (NH4)OH was added to the solution with constant stirring. The resulting molybdate / tungstate solution was stirred for 10 minutes and the pH was monitored. The solution had a pH in the range of 9-10. A second solution containing 174.65 g of Ni(NO3)2·6H2O dissolved in 150 ml of deionized water was prepared and heated to 90°C. The hot nickel solution was then slowly added to the molybdate / tungstate solution over 1 hour. The resulting mixture was heated to 91°C and stirring was continued for 30 minutes. The pH of the solution was in the range of 5-6. A blue-green precipitate formed and was collected by filtration. The precipitate was dispersed in a solution of 10.54 g of maleic acid dissolved in 1.8 L of DI water and heated to 70°C. The resulting slurry was stirred at 70°C for 30 minutes, filtered, and the collected precipitate was dried under vacuum at room temperature overnight. The material was then further dried at 120°C for 12 hours. The resulting material had a typical XRD pattern with a broad peak at 2.5 Å, indicating an amorphous Ni-OH-containing material. The BET surface area of ​​the resulting material was 101 m. 2 / g, the average pore volume was about 0.12 to 0.14 cc / g, and the average pore diameter was about 5 nm.

[0118] [Example 2] Co-Mo-W-maleate catalyst precursor Formula (NH4){[Co 3.0 (OH) 3.0-c (C4H2O4 2- ) c / 2 ](Mo 0.34 W 0.66 A catalyst precursor for (O4)2 was prepared as follows: 2.0 g of maleic acid was dissolved in 800 g of deionized water at room temperature. The pH of the resulting solution was in the range of 2-3. 17.65 g of ammonium heptamolybdate (NH4)6Mo7O 24 After dissolving 4H2O powder in the above solution, 24.67g of ammonium metatungstate (NH4)6H2W 12 O40 xH2O (>66.5% W) was added. The pH of the resulting solution was in the range of 4-5. With constant stirring, 30 ml of concentrated (NH4)OH was added to the solution. The resulting molybdate / tungstart solution was stirred for 10 minutes and the pH was monitored. The solution had a pH in the range of 9-10 at room temperature and was heated to 90°C. A second solution containing 58.28 g of cobalt nitrate dissolved in 50 g of deionized water was then prepared. The hot cobalt solution was then slowly added to the hot molybdate / tungstart solution over a 25-minute period. The resulting mixture was continuously stirred at 90°C for 1 hour. The pH of the solution was approximately 6. The dark purple-brown precipitate formed during this process was collected by filtration. The precipitate was dispersed in 250 g of DI water at 70°C. The resulting slurry was stirred for 30 minutes, filtered, and the collected precipitate was vacuum dried at room temperature overnight. The material was then further dried at 120°C for 12 hours.

[0119] [Example 3] Co-Mo-W catalyst precursor Formula (NH4) + {[Co 3.31 (OH) 3.62 ](Mo 0.3 W 0.7 The catalyst precursor of {O4)2} was prepared according to the following procedure: 17.65 g of ammonium heptamolybdate (NH4)6Mo7O 24 4H2O powder was dissolved in 800.00 g of deionized water at room temperature, followed by 24.66 g of ammonium metatungstate (NH4)6H2W. 12 O 40·xH2O (>66.5% W) was added. The pH of the resulting solution was in the range of 5.2–5.4. A second solution containing 58.26 g of cobalt nitrate hexahydrate dissolved in 50.0 g of deionized water was prepared. The pH of the resulting solution was in the range of 1–2. With constant stirring, 30 ml of concentrated (NH4)OH was added to the solution. Initially, a moss-green precipitate formed, which then changed to a two-phase mixture with a greenish suspension at the bottom and a brownish upper layer. The cobalt-containing mixture was then slowly added to the molybdate / tungstate solution over a 25-minute period at room temperature. The pH of the resulting solution was in the range of 8–8.5. The mixture was heated to 80°C and continuously stirred for 1 hour. The purplish-gray suspension was filtered while hot. The precipitate was dispersed in 2.5 L of DI water at 70°C. The resulting slurry was stirred for 30 minutes (pH 7.6), filtered, and the collected precipitate was dried under vacuum at room temperature overnight. The material was then further dried at 120°C for 12 hours.

[0120] [Example 4] Extrusion process In this example, 40 g of dried catalyst precursor prepared according to Examples 1-3 was mixed with 0.8 g of METHOCEL™ (a methylcellulose and hydroxypropyl methylcellulose polymer commercially available from DuPont) and approximately 7 g of DI water was added. An additional 7 g of water was slowly added until the mixture reached an extrudable consistency. The mixture was then extruded and dried at 120°C under N2 before sulfiding.

[0121] [Example 5] Sulfurized DMDS liquid phase The catalyst precursors of Examples 1 to 3 were placed in a tubular reactor. 2(g) bottom, 8ft 3 The temperature was increased from room temperature to 250°F at a rate of 100°F / hour. The reaction was continued for 1 hour, after which N2 was switched off and 8 ft 3The catalyst precursor was then heated to 430°F at a rate of 25°F / hr, and the temperature was exchanged with H2 at 200 psig for 1 hour. Light VGO oil (endpoint below 950°F) was pumped over the catalyst precursor at 250°F at a rate of 130 cc / hr (1 LHSV) while maintaining a hydrogen gas rate of 8 cubic feet per hour. The catalyst precursor was then heated to 430°F at a rate of 25°F / hr, and dimethyl disulfide (DMDS) was added to the light VGO at a rate of 4 cc / hr for approximately 4 hours. The catalyst precursor was then heated to 600°F, and the rate of DMDS addition was increased to 8 cc / hr. The temperature was maintained at 600°F for 2 hours, after which sulfiding was complete.

[0122] [Example 6] Sulfurization with DMDS gas phase The catalyst precursors of Examples 1 to 3 extruded according to Example 4 were placed in a tubular reactor. 2(g) bottom, 8ft 3 The temperature was increased at a rate of 100°F / hr to 450°F. The reaction was continued for 1 hour, after which N2 was switched off and 8 ft 3 The catalyst precursor was then heated to 600°F and the rate of DMDS addition was increased to 8 cc / hr. The temperature was maintained at 600°F for 2 hours, after which sulfidation was completed.

Claims

1. 1. A hydrocracking process for converting a petroleum feedstock into lower boiling point products using a two-stage hydrocracking reactor system, comprising: the reactor system includes a first zone, a pretreatment zone, a second zone, a mixed metal sulfide (MMS) catalyst zone following the pretreatment zone, and a third zone, a hydrocracking zone; The hydrocracking method comprises: (i) hydrotreating a petroleum feedstock, which is a vacuum gas oil boiling up to 593°C or a distillate fraction boiling between 175°C and 375°C, in the presence of hydrogen in a pretreatment zone over a catalyst comprising at least one Group VIII metal and at least one Group VIB metal to produce a hydrotreated effluent stream comprising a liquid product; (ii) passing the hydrotreated effluent stream through an MMS catalyst zone containing an MMS catalyst for reaction to produce a resultant effluent, wherein the MMS catalyst has the general formula: A v [(M P ) (OH) x (L) ny ] z (M VIB O 4 ) (In the formula, (v-2+2z-xxz+nxyxz) is 0, A is one monovalent cationic species; M P is a promoter metal having an oxidation state of +2 or +4 selected from one or more of Group IIA, Group IIB, Group IVA, and Group VIII metals; L is an organic oxygen-containing ligand, and M VIB is a Group VIB metal a self-supported multimetallic catalyst prepared by sulfiding a precursor catalyst in hydroxide or oxide form of (iii) passing at least a portion of the resulting effluent from (ii) through a hydrocracking zone containing a nickel, nickel-cobalt-molybdenum, cobalt-molybdenum, nickel-tungsten, or nickel-molybdenum catalyst to react to produce a hydrocracked effluent stream; Including, the power of said pretreatment zone (i) is maintained at a level of at least 56%, said power being the temperature rise exhibited by said pretreatment zone relative to the temperature rise in all three zones; A hydrocracking process in which the power is controlled by determining the temperature increase or change in each zone, and then adding more H2 to quench or more heat to increase activity as needed to maintain the power at a level of at least 56%.

2. 10. The method of claim 1, wherein the MMS catalyst in the MMS catalyst zone comprises a self-supported multimetallic catalyst prepared from precursors in hydroxide form.

3. 10. The process of claim 1, wherein the hydrocracked effluent is sent to a distillation column.

4. 10. The process of claim 1, wherein the hydrocracking zone comprises up to three reaction zones.

5. 10. The process of claim 1, wherein the hydrocracking zone comprises at least one reaction zone that is a hydrodesulfurization zone.

6. The method described in claim 5, wherein the hydrocracking zone includes a bottom reaction zone, which is a hydrodesulfurization zone, and the effluent from the hydrodesulfurization zone is sent to a distillation column.

7. 4. The method of claim 3, wherein a portion of the bottoms from the distillation column is sent to an FCC unit.

8. 4. The process of claim 3, wherein light naphtha, heavy naphtha, kerosene, and diesel fractions are recovered from the distillation column.

9. The catalyst precursor M Ｐ : M ＶＩＢ The method of claim 1 , wherein has an atomic ratio of 100:1 to 1:

100.

10. M Ｐ is nickel (Ni), and M ＶＩＢ 2. The method of claim 1, wherein is selected from molybdenum (Mo), tungsten (W), or a combination thereof.

11. The method of claim 10, wherein the catalyst precursor comprises Ni—Mo—W.

12. 12. The method of claim 11, wherein Ni:(Mo+W) has a molar ratio of 10:1 to 1:

10.

13. The method of claim 1 , wherein L is a maleate ligand.

14. 10. The process of claim 1, wherein the pretreatment zone, the MMS catalyst zone, and the hydrocracking zone are in the same reactor.

15. 10. The process of claim 1, wherein the pretreatment zone is in a first reactor and the MMS catalyst zone and the hydrocracking zone are in a second reactor.

16. 16. The method of claim 15, wherein the MMS catalyst zone is in the top half of the second reactor and comprises a catalyst prepared from a precursor in hydroxide form.

17. 10. The process of claim 1, wherein the pretreatment zone in (i) comprises a plurality of hydrotreating reaction beds prior to the MMS catalyst zone and the hydrocracking zone.

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