Classifications

• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
  • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
  • C10G47/10—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used with catalysts deposited on a carrier
  • C10G47/12—Inorganic carriers
  • C10G47/16—Crystalline alumino-silicate carriers
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J29/00—Catalysts comprising molecular sieves
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
  • B01J23/16—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
  • B01J23/24—Chromium, molybdenum or tungsten
  • B01J23/28—Molybdenum
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
  • B01J23/16—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of arsenic, antimony, bismuth, vanadium, niobium, tantalum, polonium, chromium, molybdenum, tungsten, manganese, technetium or rhenium
  • B01J23/24—Chromium, molybdenum or tungsten
  • B01J23/30—Tungsten
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
  • B01J23/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper
  • B01J23/74—Iron group metals
  • B01J23/75—Cobalt
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
  • B01J23/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper
  • B01J23/74—Iron group metals
  • B01J23/755—Nickel
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J29/00—Catalysts comprising molecular sieves
  • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J29/00—Catalysts comprising molecular sieves
  • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
  • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
  • B01J29/08—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the faujasite type, e.g. type X or Y
  • B01J29/084—Y-type faujasite
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J29/00—Catalysts comprising molecular sieves
  • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
  • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
  • B01J29/70—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of types characterised by their specific structure not provided for in groups B01J29/08 - B01J29/65
  • B01J29/7007—Zeolite Beta
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J29/00—Catalysts comprising molecular sieves
  • B01J29/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites
  • B01J29/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
  • B01J29/80—Mixtures of different zeolites
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J35/00—Catalysts, in general, characterised by their form or physical properties
  • B01J35/60—Catalysts, in general, characterised by their form or physical properties characterised by their surface properties or porosity
  • B01J35/64—Pore diameter
  • B01J35/653—500-1000 nm
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
  • C10G45/02—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing
  • C10G45/04—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used
  • C10G45/12—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used containing crystalline alumino-silicates, e.g. molecular sieves
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
  • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
  • C10G47/02—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used
  • C10G47/10—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions characterised by the catalyst used with catalysts deposited on a carrier
  • C10G47/12—Inorganic carriers
  • C10G47/16—Crystalline alumino-silicate carriers
  • C10G47/20—Crystalline alumino-silicate carriers the catalyst containing other metals or compounds thereof

Definitions

• the invention relates to catalyst compositions and their use in hydrocarbon conversion processes, particularly hydrocracking.
• the invention more specifically relates to a catalyst composition that comprises a Y zeolite and a beta zeolite as active cracking components.
• the invention specifically relates to a hydrocracking process that produces middle distillate.
• Petroleum refiners often produce desirable products such as turbine fuel, diesel fuel, and other hydrocarbon liquids known as middle distillates, as well as lower boiling liquids such as naphtha and gasoline, by hydrocracking a hydrocarbon feedstock derived from crude oil. Hydrocracking also has other beneficial results such as removing sulfur and nitrogen from the feedstock by hydro treating. Feedstocks most often subject to hydrocracking are gas oils and heavy gas oils recovered from crude oil by distillation.
• Hydrocracking is generally carried out by contacting, in an appropriate reactor vessel, the gas oil or other hydrocarbon feedstock with a suitable hydrocracking catalyst under appropriate conditions, including an elevated temperature and an elevated pressure and the presence of hydrogen so as to yield a lower overall average boiling point product containing a distribution of hydrocarbon products desired by the refiner. Although the operating conditions within a hydrocracking reactor have some influence on the yield of the products, the hydrocracking catalyst is a prime factor in determining such yields. [0004] Hydrocracking catalysts are subject to initial classification on the basis of the nature of the predominant cracking component of the catalyst.
• hydrocracking catalysts are also subject to classification on the basis of their intended predominant product of which the two main products are naphtha and "distillate", a term which in the hydrocracking refining art refers to distillable petroleum derived fractions having a boiling point range that is above that of naphtha.
• Distillate typically includes the products recovered at a refinery as kerosene and diesel fuel. At the present time, distillate is in high demand. For this reason, refiners have been focusing on hydrocracking catalysts which selectively produce a distillate fraction.
• the three main catalytic properties by which the performance of a hydrocracking catalyst for making distillate is evaluated are activity, selectivity, and stability. Activity may be determined by comparing the temperature at which various catalysts must be utilized under otherwise constant hydrocracking conditions with the same feedstock so as to produce a given percentage, normally 65 percent, of products boiling in the desired range, e.g., below 371°C (700 0 F) for distillate. The lower the temperature required for a given catalyst, the more active such a catalyst is in relation to a catalyst requiring a higher temperature.
• Selectivity of hydrocracking catalysts may be determined during the foregoing described activity test and is measured as a percentage of the fraction of the product boiling in the desired distillate product range, e.g., from 149°C (300 0 F) to 371°C (700 0 F).
• Stability is a measure of how well a catalyst maintains its activity over an extended time period when treating a given hydrocarbon feedstock under the conditions of the activity test. Stability is generally measured in terms of the change in temperature required per day to maintain a 65 percent or other given conversion.
• hydrocracking catalysts containing a Y zeolite having a unit cell size or dimension a o of from 24.25 to 24.32 angstrom and containing a beta zeolite preferably having an overall silica to alumina (Si ⁇ 2 to AI2O3) mole ratio of less than 30 and a SFg adsorption capacity of at least 28 weight-percent (hereinafter wt-%) have substantially improved selectivity at a given activity and substantially improved activity at a given selectivity compared to other hydrocracking catalysts now commercially available for use in hydrocracking processes for producing distillate.
• the catalyst also contains a metal hydrogenation component such as nickel, cobalt, tungsten, molybdenum, or any combination thereof.
• the catalyst contains from more than 2 wt-% to at most 5 wt-% beta zeolite based on the combined weight of the beta zeolite, the Y zeolite, and the support on a dried basis, and the catalyst has a weight ratio of the Y zeolite to the beta zeolite of from 2.5 to 12.5 on a dried basis.
• the Y zeolite has an overall silica to alumina mole ratio of from 5.0 to 11.0.
• the catalyst contains no additional zeolites or no additional beta zeolite or Y zeolite.
• the catalyst consists of, or consists essentially of, a hydrogenation component, a beta zeolite having an overall silica to alumina mole ratio of less than 30 and a SF5 adsorption capacity of at least 28 wt-%, a Y zeolite having a unit cell size from 24.25 to 24.32 angstrom, and a support, wherein the Y zeolite has an overall silica to alumina mole ratio of from 5.0 to 11.0, wherein the catalyst contains from more than 2 wt-% to at most 5 wt-% beta zeolite based on the combined weight of the beta zeolite, the Y zeolite, and the support on a dried basis, and wherein the catalyst has a weight ratio of the Y zeolite to the beta zeolite of from 2.5 to 12.5 on a dried basis, wherein the Y zeolite has a surface area of less than 800 m 2 /g.
• the product contains a relatively large proportion of components boiling in the distillate range, which as defined herein is from 149 0 C (300 0 F) to 371 0 C (700 0 F).
• Beta and Y zeolites have been proposed in combination as components of several different catalysts including catalysts for hydrocracking.
• catalysts for hydrocracking For instance, US-A-5,275,720; US- A-5,279,726; and US-A-5,35O,5O1 describe hydrocracking processes using a catalyst comprising a beta zeolite and a Y zeolite US-A-5,275,720 describes a hydrocracking process using a catalyst comprising, among other components, zeolite beta and a Y zeolite having a unit cell size between 24.25 and 24.35 angstrom and a water vapor sorption capacity at 4.6 mm water vapor partial pressure and 25 0 C less than 8.0 percent by weight of the zeolite.
• US- Al-2004/0152587 describes a hydrocracking catalyst comprising a carrier comprising a zeolite of the faujasite structure having a unit cell size in the range of from 24.10 to 24.40 angstrom, a bulk silica to alumina ratio above 12, and a surface area of at least 850 m ⁇ /g, and the catalyst may contain a second zeolite such as beta zeolite, ZSM-5 zeolite, or a Y zeolite of a different unit cell size.
• Two different Y zeolites have also been proposed in combination as components of several different catalysts including catalysts for hydrocracking, as described in US-A-4,661,239 and US-A-4,925,546.
• the process and composition disclosed herein may be used to convert a feedstock containing organic compounds into products, particularly by acid catalysis, such as hydrocracking organic compounds especially hydrocarbons into a product of lower average boiling point and lower average molecular weight.
• the composition which may be a catalyst and/or a catalyst support, comprises a beta zeolite and a Y zeolite.
• the composition may also comprise a refractory inorganic oxide.
• the composition When used as a catalyst for hydrocracking, the composition contains a beta zeolite, Y zeolite, a refractory inorganic oxide, and a hydrogenation component.
• the hydrocracking process and composition disclosed herein centers on using a catalyst containing a particular beta zeolite and a particular Y zeolite.
• the composition preferably contains no additional zeolites, no additional beta zeolite, and/or no additional Y zeolite.
• the beta zeolite preferably has a relatively low silica to alumina mole ratio and a relatively high SF5 adsorption capacity.
• the Y zeolite has a unit cell size of from 24.25 to
• Beta zeolite is well known in the art as a component of hydrocracking catalysts. Beta zeolite is described in US-A-3,308,069 and US Reissue No. 28341, which are hereby incorporated by reference herein in their entireties.
• the beta zeolite that is used in the process and composition disclosed herein has a silica to alumina mole ratio of less than 30 in one embodiment, less than 25 in another embodiment, more than 9 and less than 30 in yet another embodiment, more than 9 and less than 25 in a further embodiment, more than 20 and less than 30 in another embodiment, or more than 15 and less than 25 in still another embodiment.
• the silica to alumina (Si ⁇ 2 to AI2O3) mole ratio of a zeolite is the mole ratio as determined on the basis of the total or overall amount of aluminum and silicon (framework and non-framework) present in the zeolite, and is sometimes referred to herein as the overall silica to alumina (Si ⁇ 2 to AI2O3) mole ratio.
• Beta zeolite is usually synthesized from a reaction mixture containing a templating agent.
• templating agents for synthesizing beta zeolite is well known in the art.
• US-A-3, 308,069 and US Reissue No. 28341 describe using tetraethylammonium hydroxide and US-A-5, 139,759, which is hereby incorporated herein by reference in its entirety, describes using the tetraethylammonium ion derived from the corresponding tetraethylammonium halide.
• Another standard method of preparing beta zeolite is described in the book titled Verified Synthesis of Zeolitic Materials, by H.
• the beta zeolite is calcined in air at a temperature of from 500 to 700 0 C (932 to 1292°F) for a time sufficient to remove to remove the templating agent from the beta zeolite. Calcination to remove the templating agent can be done before or after the beta zeolite is combined with the support and/or the hydrogenation component.
• the templating agent could be removed at calcination temperatures above 700 0 C (1292°F), very high calcination temperatures could significantly decrease the SFg adsorption capacity of beta zeolite. For this reason it is believed that calcination temperatures above 750 0 C (1382°F) for removing the templating agent should be avoided when preparing the beta zeolite for use in the process disclosed herein. It is critical to the process disclosed herein that the SFg adsorption capacity of the beta zeolite is at least 28 wt-%.
• the SF5 adsorption capacity of the beta zeolite should be at least 28 wt-%.
• the beta zeolite of the process and composition disclosed herein may be characterized in terms of SF5 adsorption.
• SF5 adsorption This is a recognized technique for the characterization of microporous materials such as zeolites. It is similar to other adsorption capacity measurements, such as water capacity, in that it uses weight differences to measure the amount of SF5 which is adsorbed by a sample which has been pretreated to be substantially free of the adsorbate. SF5 is used in this test since because its size and shape hinders its entrance into pores having a diameter of less than 6 angstrom. It thus can be used as one measurement of available pore mouth and pore diameter shrinkage. This in turn is a measurement of the effect of steaming on the zeolite.
• the sample is preferably first predried in a vacuum at 300 0 C (572°F) for one hour, then heated at atmospheric pressure in air at 650 0 C (1202 0 F) for two hours, and finally weighed. It is then exposed to the SF5 for one hour while the sample is maintained at a temperature of 20 0 C (68 0 F). The vapor pressure of the SF5 is maintained at that provided by liquid SF5 at 400 torr (53.3 kPa (7.7 psi)). The sample is again weighed to measure the amount of adsorbed SF5. The sample may be suspended on a scale during these steps to facilitate these steps.
• any mass production procedure involving techniques such as steaming and heating there is a possibility for individual particles to be subjected to differing levels of treatment. For instance, particles on the bottom of a pile moving along a rotary kiln may not be subjected to the same atmosphere or temperature as the particles which cover the top of the pile. This factor must be considered during manufacturing and also during analysis and testing of the finished product. It is, therefore, recommended that any test measure done on the material is performed on a representative composite sample of the entire quantity of finished product to avoid being misled by measurements performed on individual particles or on a non-representative sample. For instance, an adsorption capacity measurement is made on a representative composite sample.
• the process and the composition disclosed herein can use a beta zeolite that has not been subjected to a steaming treatment
• the process and the composition disclosed herein can also use beta zeolite that is subjected to steaming, provided that the steaming is relatively mild in comparison to steaming of beta zeolite in the literature.
• steaming beta zeolite can yield a catalyst that can be used in the process and composition disclosed herein.
• Hydrothermally treating zeolites for use in hydrocracking catalysts is a relatively blunt tool. For any given zeolite, steaming decreases the acidity of the zeolite.
• the beta zeolite is to be steamed, such steaming can be performed successfully in different ways, with the method which is actually employed commercially often being greatly influenced and perhaps dictated by the type and capability of the available equipment. Steaming can be performed with the beta zeolite retained as a fixed mass or with the beta zeolite being confined in a vessel or being tumbled while confined in a rotating kiln. The important factors are uniform treatment of all beta zeolite particles under appropriate conditions of time, temperature and steam concentration. For instance, the beta zeolite should not be placed such that there is a significant difference in the amount of steam contacting the surface and the interior of the beta zeolite mass.
• the beta zeolite may be steam treated in an atmosphere having live steam passing through the equipment providing low steam concentration. This may be described as being at a steam concentration of a positive amount less than 50 mol-%. Steam concentrations may range from 1 to 20 mol-% or from 5 to 10 mol-%, with small-scale laboratory operations extending toward higher concentrations. The steaming may be performed for a positive time period of less than or equal to 1 or 2 hours or for 1 to 2 hours at a temperature of less than or equal to 600 0 C (1112°F) at atmospheric pressure and a positive content of steam of less than or equal to 5 mol-%.
• the steaming may be performed for a positive time period of less than or equal to 2 hours at a temperature of less than or equal to 650 0 C (1202 0 F) at atmospheric pressure and a positive content of steam of less than or equal to 10 mol-%.
• the steam contents are based on the weight of vapors contacting the beta zeolite. Steaming at temperatures above 650 0 C (1202 0 F) appears to result in beta zeolite that is not useful in the process disclosed herein since the SFg adsorption capacity of the resulting beta zeolite is too low.
• the steaming temperature can be from 600 0 C (1112°F) to 650 0 C (1202 0 F), or less than 600 0 C (1112°F). It is taught in the art that there is normally an interplay between time and temperature of steaming, with an increase in temperature reducing the required time. Nevertheless, if steaming is done, for good results it appears a time period of Vi to 2 hours or 1 to 1 Vi hours can be used.
• the method of performing steaming on a commercial scale may be by means of a rotary kiln having steam injected at a rate which maintains an atmosphere of 10 mol-% steam.
• An exemplary lab scale steaming procedure is performed with the zeolite held in a 6.4 cm (2-1/2 inch) quartz tube in a clam shell furnace. The temperature of the furnace is slowly ramped up by a controller. After the temperature of the zeolite reaches 150 0 C (302 0 F) steam generated from deionized water held in a flask is allowed to enter the bottom of the quartz tube and pass upward. Other gas can be passed into the tube to achieve the desired steam content. The flask is refilled as needed. In the exemplary procedure the time between cutting in the steam and the zeolite reaching 600 0 C (1112°F) is one hour.
• the temperature in the furnace is reduced by resetting the controller to 20 0 C (68°F).
• the furnace is allowed to cool to 400 0 C (752°F) (about 2 hours) and the flow of steam into the quartz tube is stopped.
• the sample is removed at 100 0 C (212°F) and placed in a lab oven held overnight at 110 0 C (230 0 F) with an air purge.
• the beta zeolite of the process and composition disclosed herein is not treated with an acid solution to effect dealumination.
• essentially all raw (as synthesized) beta zeolite is exposed to an acid to reduce the concentration of alkali metal (e.g., sodium) which remains from synthesis.
• alkali metal e.g., sodium
• This step in the beta zeolite manufacture procedure is not considered part of the treatment of manufactured beta zeolite as described herein.
• the beta zeolite is exposed to an acid only during incidental manufacturing activities such as peptization during forming or during metals impregnation.
• the beta zeolite is not acid washed after the steaming procedure as to remove aluminum "debris" from the pores.
• a Y zeolite having a unit cell size of from 24.25 to 24.32 angstrom.
• the Y zeolite preferably has a unit cell size of from 24.26 to 24.30 angstrom.
• the Y zeolite can have an overall silica to alumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from 5.0 to 10.0 in yet another embodiment.
• the process and composition disclosed herein require a Y zeolite.
• Y zeolite as used herein is meant to encompass all crystalline zeolites having either the essential X-ray powder diffraction pattern set forth in US-A-3, 130,007 or a modified Y zeolite having an X-ray powder diffraction pattern similar to that of US-A- 3,130,007 but with the d-spacings shifted somewhat due, as those skilled in the art will realize, to cation exchanges, calcinations, etc., which are generally necessary to convert the Y zeolite into a catalytically active and stable form.
• the Y zeolites are modified Y zeolites in comparison to the Y zeolite taught in US-A-3, 130,007.
• unit cell size means the unit cell size as determined by X-ray powder diffraction.
• the Y zeolites used in the process and composition disclosed herein are large pore zeolites having an effective pore size greater than 7.0 angstrom. Since some of the pores of the Y zeolites are relatively large, the Y zeolites allow molecules relatively free access to their internal structure. The pores of the Y zeolites permit the passage thereinto of benzene molecules and larger molecules and the passage therefrom of reaction products.
• Y zeolites that may be used in the process and composition disclosed herein includes zeolites that are sometimes referred to as ultrastable or ultrahydrophobic Y zeolites.
• the composition and properties of this group of Y zeolites are, in essence, prepared by a four step procedure. First, a Y zeolite in the alkali metal form (usually sodium) and typically having a unit cell size of 24.65 angstrom is cation exchanged with ammonium ions.
• the ammonium exchange step typically reduces the sodium content of the starting sodium Y zeolite from a value usually greater than 8 wt-%, usually from 10 to 13 wt-%, calculated as Na2 ⁇ , to a value in the range from 0.6 to 5 wt-%, calculated as Na2 ⁇ .
• Methods of carrying out the ion exchange are well known in the art.
• the Y zeolite from the first step is calcined in the presence of water vapor.
• the Y zeolite is calcined in the presence of at least 1.4 kPa(absolute) (hereinafter kPa(a)) (0.2 psi(absolute) (hereinafter psi(a))), at least 6.9 kPa(a) (1.0 psi(a)), or at least 69 kPa(a) (10 psi(a)) water vapor, in three embodiments.
• the Y zeolite is calcined in an atmosphere consisting essentially of or consisting of steam.
• the Y zeolite is calcined so as to produce a unit cell size in the range of 24.40 to 24.64 angstrom.
• the Y zeolite from the second step is ammonium exchanged once again.
• the second ammonium exchange further reduces the sodium content to less than 0.5 wt-%, usually less than 0.3 wt-%, calculated as Na2 ⁇ .
• the Y zeolite from the third step is treated further so as to yield Y zeolite having a unit cell size from 24.25 to 24.32 angstrom or preferably from 24.26 to 24.30 angstrom.
• the zeolite Y resulting from the fourth step has an overall silica to alumina mole ratio from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from 5.0 to 10.0 in yet another embodiment.
• the treatment of the fourth step can comprise any of the well known techniques for dealuminating zeolites in general and ultrastable Y zeolite in particular so as to yield the desired unit cell size and overall silica to alumina mole ratio.
• the fourth treatment step may change the unit cell size and/or the framework silica to alumina mole ratio, with or without changing the overall silica to alumina mole ratio.
• zeolite dealumination is accomplished by chemical methods such as treatments with acids, e.g., HCl, with volatile halides, e.g., SiCl ⁇ or with chelating agents such as ethylenediaminetetraacetic acid (EDTA).
• acids e.g., HCl
• volatile halides e.g., SiCl ⁇
• chelating agents such as ethylenediaminetetraacetic acid (EDTA).
• Another common technique is a hydrothermal treatment of the zeolite in either pure steam or in air/steam mixtures, preferably such as calcining in the presence of sufficient water vapor (for example, in an atmosphere consisting essentially of steam, and most preferably consisting of steam) so as to yield the desired unit cell size and overall silica to alumina mole ratio.
• sufficient water vapor for example, in an atmosphere consisting essentially of steam, and most preferably consisting of steam
• US-A-3 ,929,672 which is hereby incorporated herein by reference in its entirety, discloses a method for dealuminating an ultrastable Y zeolite.
• US- A-3, 929,672 teaches a preparation procedure wherein a sodium Y zeolite is partially exchanged with ammonium ions, followed by steam calcination under controlled temperature and steam partial pressure, followed by yet another ammonia exchange and then by an optional calcination step in a dry atmosphere. The exchange and steam calcination steps can be repeated to achieve the desired degree of dealumination and unit cell size reduction.
• zeolites of US- A-3, 929,672 are known under the designation Y- 84 or LZY-84 commercially available from UOP LLC, Des Plaines, Illinois, U.S.A.
• Y-84 or LZY-84 zeolites may be produced by the first three steps just mentioned, but optionally one may include a further calcination step in a dry atmosphere, e.g., a calcination in water- and steam-free air, at 482°C (900 0 F) or higher.
• the Y zeolite produced by the four-step procedure in US-A-5,350,501 is a UHP- Y zeolite, an ultrahydrophobic Y zeolite as defined in US-A-5,350,501.
• US-A-5,350,501 defines a "UHP-Y" zeolites as zeolite aluminosilicates having among other properties, a unit cell size or dimension as of less than 24.45 angstrom and a sorptive capacity for water vapor at 25°C and a p/p 0 value of 0.10 of less than 10.00 weight percent.
• the most preferred UHP- Y zeolite in US-A-5,350,501 is LZ-10.
• Y zeolites which may be used in the process and composition disclosed herein may be prepared by dealuminating a Y zeolite having an overall silica to alumina mole ratio below 5 and are described in detail in US-A-4,503,023; US-A-4,597,956 and US-A-4,735,928; which are hereby incorporated herein by reference in their entireties.
• US-A-4,503,023 discloses another procedure for dealuminating a Y zeolite involving contacting the Y zeolite with an aqueous solution of a fluorosilicate salt using controlled proportions, temperatures, and pH conditions which avoid aluminum extraction without silicon substitution.
• the fluorosilicate salt is used as the aluminum extractant and also as the source of extraneous silicon which is inserted into the Y zeolite structure in place of the extracted aluminum.
• the salts have the general formula: (A) 2 Zb SiF 6 wherein A is a metallic or nonmetallic cation other than H + having the valence "b.”
• Cations represented by "A” are alkylammonium, NH4+, Mg + +, Li + , Na + , K + , Ba + +, Cd + +, Cu + +, H + , Ca + +, Cs + , Fe + +, Co + +, Pb + +, Mn + +, Rb + , Ag + , Sr + +, Ti + , and Zn + +.
• LZ-210 a zeolitic aluminosilicate molecular sieve commercially available from UOP LLC, Des Plaines, Illinois, U.S.A.
• LZ-210 zeolites and the other zeolites of this group are conveniently prepared from a Y zeolite starting material.
• the LZ-210 zeolite has an overall silica to alumina mole ratio from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from 5.0 to 10.0 in yet another embodiment.
• the unit cell size can be from 24.25 to 24.32 angstrom or preferably from 24.26 to 24.30 angstrom in the case of Y Zeolite I.
• the unit cell size can be from 24.33 to 24.38 angstrom or preferably from 24.34 to 24.36 angstrom.
• the LZ-210 class of zeolites used in the process and composition disclosed herein have a composition expressed in terms of mole ratios of oxides as in the following formula:
• LZ-210 zeolites may be prepared by dealuminating Y-type zeolites using an aqueous solution of a fluorosilicate salt, preferably a solution of ammonium hexafluorosilicate.
• the dealumination can be accomplished by placing a Y zeolite, normally but not necessarily an ammonium exchanged Y zeolite, into an aqueous reaction medium such as an aqueous solution of ammonium acetate, and slowly adding an aqueous solution of ammonium fluorosilicate. After the reaction is allowed to proceed, a zeolite having an increased overall silica to alumina mole ratio is produced.
• the magnitude of the increase is dependent at least in part on the amount of fluorosilicate solution contacted with the zeolite and on the reaction time allowed. Normally, a reaction time of between 10 and 24 hours is sufficient for equilibrium to be achieved.
• the resulting solid product which can be separated from the aqueous reaction medium by conventional filtration techniques, is a form of LZ-210 zeolite. In some cases this product may be subjected to a steam calcination by methods well known in the art.
• the product may be contacted with water vapor at a partial pressure of at least 1.4 kPa(a) (0.2 psi(a)) for a period of between 1/4 to 3 hours at a temperature between 482°C (900 0 F) and 816°C (1500 0 F) in order to provide greater crystalline stability.
• the product of the steam calcination may be subjected to an ammonium-exchange by methods well known in the art.
• the product may be slurried with water after which an ammonium salt is added to the slurry. The resulting mixture is typically heated for a period of hours, filtered, and washed with water.
• Methods of steaming and ammonium-exchanging LZ-210 zeolite are described in US-A-4,503,023; US- A-4,735,928 and US-A-5,275,720.
• Y zeolites prepared by the above-discussed preparation procedures and used in the process and composition disclosed herein have the essential X-ray powder diffraction pattern of zeolite Y and a unit cell size or dimension a 0 of from 24.25 to 24.32 angstrom, preferably from 24.26 to 24.30 angstrom.
• the Y zeolite can have an overall silica to alumina mole ratio of from 5.0 to 12.0 in one embodiment, from 5.0 to 11.0 in another embodiment, and from 5.0 to 10.0 in yet another embodiment.
• the Y zeolite may have a surface area (BET) of at least 500 m 2 /g, less than 800 m 2 /g, often less than 700 m 2 /g and typically from 500 to 650 m 2 /g.
• BET surface area
• Another method of increasing the stability and/or acidity of the Y zeolites is by exchanging the Y zeolite with polyvalent metal cations, such as rare earth-containing cations, magnesium cations or calcium cations, or a combination of ammonium ions and polyvalent metal cations, thereby lowering the sodium content until it is as low as the values described above after the first or second ammonium exchange steps. Methods of carrying out the ion exchange are well known in the art.
• the catalyst used in the process disclosed herein is intended primarily for use as a replacement catalyst in existing commercial hydrocracking units. Its size and shape is, therefore, preferably similar to those of conventional commercial catalysts. It is preferably manufactured in the form of a cylindrical extrudate having a diameter of from 0.8 -3.2 mm (1/32 - 1/8 in). The catalyst can however be made in any other desired form such as a sphere or pellet. The extrudate may be in forms other than a cylinder such as the form of a well- known trilobal or other shape which has advantages in terms or reduced diffusional distance or pressure drop.
• the catalyst contains a number of non-zeolitic materials. This is for several reasons such as particle strength, cost, porosity, and performance.
• the other catalyst components therefore, make positive contributions to the overall catalyst even if not as active cracking components.
• These other components are referred to herein as the support.
• Some traditional components of the support such as silica-alumina normally make some contribution to the cracking capability of the catalyst.
• the catalyst contains a relatively small content of beta zeolite.
• the catalyst contains more than 2 wt-% to at most 5 wt-%, preferably more than 2 wt- % to at most 3 wt-%, of beta zeolite based on the combined weight of the beta zeolite, Y zeolite, and the support all on a dried basis.
• the weight on a dried basis is considered to be the weight after heating in dry air at 500 0 C (932°F) for 6 hours.
• the catalyst has a weight ratio of the Y zeolite to the beta zeolite of from 2.5 to 12.5, preferably from 2.5 to 3.0, on a dried basis.
• the remainder of the catalyst particle besides the zeolitic material may be taken up primarily by conventional hydrocracking materials such as alumina and/or silica-alumina.
• alumina and/or silica-alumina The presence of silica-alumina helps achieve the desired performance characteristics of the catalyst.
• the catalyst contains at least 25 wt-% alumina and at least 25 wt-% silica-alumina, both based on the combined weight of the zeolites and the support, all on a dried basis.
• the silica-alumina content of the catalyst is above 40 wt-% and the alumina content of the catalyst is above 20 wt-%, both based on the combined weight of the zeolites and the support, all on a dried basis.
• the alumina is believed to function only as a binder and to not be an active cracking component.
• the catalyst support may contain over 50 wt-% silica-alumina or over 50 wt-% alumina based on the weight of the support on a dried basis. Approximately equal amounts of silica-alumina and alumina are used in an embodiment.
• Other inorganic refractory materials which may be used as a support in addition to silica-alumina and alumina include for example silica, zirconia, titania, boria, and zirconia-alumina. These aforementioned support materials may be used alone or in any combination.
• the subject catalyst contains a metallic hydrogenation component.
• the hydrogenation component is preferably provided as one or more base metals uniformly distributed in the catalyst particle.
• the hydrogenation component is one or more element components from Groups 6, 9, and 10 of the periodic table.
• Noble metals such as platinum and palladium could be applied but best results have been obtained with a combination of two base metals.
• nickel or cobalt is paired with tungsten or molybdenum, respectively.
• the preferred composition of the metal hydrogenation component is both nickel and molybdenum or both nickel and tungsten.
• the amount of nickel or cobalt is preferably between 2 and 8 wt-% of the finished catalyst.
• the amount of tungsten or molybdenum is preferably between 8 and 22 wt-% of the finished catalyst.
• the total amount of a base metal hydrogenation component is from 10 to 30 wt-% of the finished catalyst.
• the catalyst of the subject process can be formulated using industry standard techniques. This can, with great generalization, be summarized as admixing the beta zeolite and the Y zeolite with the other inorganic oxide components and a liquid such as water or a mild acid to form an extrudable dough followed by extrusion through a multihole die plate.
• the extrudate is collected and preferably calcined at high temperature to harden the extrudate.
• the extruded particles are then screened for size and the hydrogenation components are added as by dip impregnation or the well known incipient wetness technique. If the catalyst contains two metals in the hydrogenation component these may be added sequentially or simultaneously.
• the catalyst particles may be calcined between metal addition steps and again after the metals are added.
• the comulling is effected with a source of metal, such as ammonium heptamolybdate or ammonium metatungstate and another source of another metal, such as nickel nitrate or cobalt nitrate, with both source compounds generally being introduced into the combined materials in the form of an aqueous solution or as a salt.
• the hydrocracking process disclosed herein will be operated within the general range of conditions now employed commercially in hydrocracking processes. The operating conditions in many instances are refinery or processing unit specific.
• the inlet temperature of the catalyst bed should be from 232°C (450 0 F) to 454°C (850 0 F), and the inlet pressure should be from 5171 kPa(g) (750 psi(g)) to 24132 kPa(g) (3500 psi(g)), and typically from 6895 kPa(g) (1000 psi(g)) to 24132 kPa(g) (3500 psi(g)).
• the feed stream is admixed with sufficient hydrogen to provide a volumetric hydrogen circulation rate per unit volume of feed of 168 to 1684 normal ltr/ltr measured at 0 0 C (32°F) and 101.3 kPa(a) (14.7 psi(a)) (1000 to 10000 standard ft 3 /barrel (SCFB) measured at 15.6°C (60 0 F) and 101.3 kPa(a) (14.7 psi(a))) and passed into one or more reactors containing fixed beds of the catalyst.
• the hydrogen will be primarily derived from a recycle gas stream which may pass through purification facilities for the removal of acid gases although this is not necessary.
• LHSV liquid hourly space velocity, which is defined as the volumetric flow rate of liquid per hour divided by the catalyst volume, where the liquid volume and the catalyst volume are in the same volumetric units.
• the typical feed to the process disclosed herein is a mixture of many different hydrocarbons and coboiling compounds recovered by fractional distillation from a crude petroleum. It will normally have components that boil higher than the upper end of the range of the 149°C (300 0 F) to 371 0 C (700 0 F) boiling range in order to produce distillate. Often it will have a boiling point range starting above 340 0 C (644°F) and ending in one embodiment below 482°C (900 0 F), in another embodiment below 540 0 C (1004 0 F), and in a third embodiment below 565°C (1049 0 F).
• Such a petroleum derived feed may be a blend of streams produced in a refinery such as atmospheric gas oil, coker gas oil, straight run gas oil, deasphalted gas oil, vacuum gas oil, and FCC cycle oil.
• a typical gas oil comprises components that boil in the range of from 166°C (330 0 F) to 566°C (1050 0 F).
• the feed to the process disclosed herein can be a single fraction such as a heavy vacuum gas oil.
• a typical heavy gas oil fraction has a substantial proportion of the hydrocarbon components, usually at least 80 percent by weight, boiling from 371°C (700 0 F) to 566°C
• Synthetic hydrocarbon mixtures such as recovered from shale oil or coal can also be processed in the subject process.
• the feed may be subjected to hydrotreating or treated as by solvent extraction prior to being passed into the subject process to remove gross amounts of sulfur, nitrogen or other contaminants such as asphalt enes.
• the subject process is expected to convert a large portion of the feed to more volatile hydrocarbons such as distillate boiling range hydrocarbons. Typical conversion rates vary from 50 to 100 volume-percent (hereinafter vol-%) depending greatly on the feed composition.
• the conversion rate is between from 60 to 90 vol-% in an embodiment of the process disclosed herein, from 70 to 90 vol-% in another embodiment, from 80 and to 90 vol- % in yet another embodiment, and from 65 to 75 vol-% in still another embodiment.
• the effluent of the process will actually contain a broad variety of hydrocarbons ranging from methane to essentially unchanged feed hydrocarbons boiling above the boiling range of any desired product.
• the effluent of the process typically passes from a reactor containing a catalyst and is usually separated by methods known to a person of ordinary skill in the art, including phase separation or distillation, to produce a product having any desired final boiling point.
• the hydrocarbons boiling above the final boiling point of any desired product are referred to as unconverted products even if their boiling point has been reduced to some extent in the process. Most unconverted hydrocarbons are recycled to the reaction zone with a small percentage, e.g. 5 wt-% being removed as a drag stream. For producing distillate, at least 30 wt-%, and preferably at least 50 wt-%, of the effluent boils below 371°C (700 0 F).
• the process and composition disclosed herein can be employed in what are referred to in the art as single stage and two stage process flows, with or without prior hydrotreating. These terms are used as defined and illustrated in the book titled Hydrocracking Science and Technology, by J. Scherzer and AJ.
• the subject catalyst can be employed in either the first or second stage.
• the catalyst may be preceded by a hydrotreating catalyst in a separate reactor or may be loaded into the same reactor as a hydrotreating catalyst or a different hydrocracking catalyst.
• An upstream hydrotreating catalyst can be employed as feed pretreatment step or to hydrotreat recycled unconverted materials.
• the hydrotreating catalyst can be employed for the specific purpose of hydrotreating polynuclear aromatic (PNA) compounds to promote their conversion in subsequent hydrocracking catalyst bed(s).
• PNA polynuclear aromatic
• the subject catalyst can also be employed in combination with a second, different catalyst, such as a catalyst based upon Y zeolite or having primarily amorphous cracking components.
• a second, different catalyst such as a catalyst based upon Y zeolite or having primarily amorphous cracking components.
• the catalyst is employed with a feed or in a configuration that the feed passing through the catalyst is a raw feed or resembles a raw feed.
• the sulfur content of crude oil, and hence the feed to this process varies greatly depending on its source.
• a raw feed is intended to refer to a feed which has not been hydrotreated or which still contains organic sulfur compounds which result in a sulfur level above 1000 wt-ppm or which still contains organic nitrogen compounds that result in a nitrogen level above 100 wt-ppm (0.01 wt-%).
• the catalyst is used with a feed that has been hydrotreated.
• a feed that has been hydrotreated Persons of ordinary skill in the art of hydrocarbon processing know and can practice hydrotreating of a raw feed to produce a hydrotreated feed to be charged to the process disclosed herein.
• the sulfur level of the hydrotreated feed may be between 500 and 1000 wt-ppm
• the sulfur level of the hydrotreated feed is less than 500 wt-ppm in one embodiment of the process disclosed herein and from 5 to 500 wt-ppm in another embodiment.
• the nitrogen level of the hydrotreated feed is less than 100 wt-ppm in one embodiment and from 1 to 100 wt-ppm in another embodiment.
• a modified Y zeolite was prepared by steaming an ammonium exchanged Y zeolite sold by UOP LLC (Des Plaines, Illinois, USA) and referred to in the literature as Y-84 having a sodium content of less than 0.2 wt-% calculated as Na2 ⁇ .
• the resulting modified Y zeolite is referred to herein as Sample 1 and had an overall silica to alumina (Si ⁇ 2 to AI2O3) mole ratio of 5.0 to 5.5, a unit cell size of 24.28 angstrom, and a surface area of 540 to 640 m 2 /g.
• Two catalysts were prepared by mixing Sample 1, a beta zeolite having an overall silica to alumina (Si ⁇ 2 to AI2O3) mole ratio of 23.8 and an SFg adsorption capacity of 29 wt-% and containing the template used during its synthesis if present, amorphous silica-alumina, and HN ⁇ 3-peptized CatapalTM C boehmite alumina in a muller.
• the amorphous silica-alumina was CCIC silica-alumina which had a nominal composition of 75 wt-% silica and 25 wt-% alumina.
• CCIC silica-alumina is available from Catalysts & Chemicals Industries Co. Ltd. (CCIC). The amounts of these components on a dried basis in each final catalyst are listed in the Table.
• the resulting mixture was extruded into 1.6 mm (1/16 in) diameter cylindrical particles of between 3.2 mm (1/8 in) and 12.7 mm (1/2 in) in length.
• the wet extrudates were dried at 104 0 C (220 0 F) for a minimum of 4 hr and then calcined at temperatures in excess of 550 0 C (1022 0 F) for a minimum of 90 minutes.
• Both of the above-described two catalysts was pre-sulfided by passing a gas stream consisting of 10 vol-% H2S and the balance H2 through a bed of the catalyst at a temperature initially of 149°C (300 0 F) and slowly raised to 413°C (775°F) and held at the temperature for 6 hours.
• the two catalysts were compared for hydrocracking activity and selectivity (i.e., product yields) in simulated first stage testing. Specifically, the two catalysts were separately tested for hydrocracking a hydrotreated light Arabian vacuum gas oil (VGO) feed having a specific gravity of 0.877 at 15.6°C (6O 0 F) (API gravity of 30.05°), an initial boiling point of 107 0 C (224°F), a 5 wt-% boiling point of 195°C (382°F), a final boiling point of 550 0 C (1021 0 F), and a 50 wt-% boiling point of 24°C (795°F), with 13 wt-% boiling below 288°C (550 0 F) and 26 wt-% boiling below 371°C (700 0 F).
• VGO hydrotreated light Arabian vacuum gas oil
• Each catalyst was tested for simulated first stage operation by passing the feedstock through a laboratory size reactor at a LHSV of 1.5 hr 1, a total pressure of 13786 kPa(g) (2000 psi(g)), and a volumetric hydrogen feed rate per unit volume of feed of 1684 normal ltr/ltr measured at 0 0 C (32°F) and 101.3 kPa(a) (14.7 psi(a)) (10000 SCFB measured at 15.6°C (6O 0 F) and 101.3 kPa(a) (14.7 psi(a))).
• Sufficient di-tert-butyl disulfide was added to the feed to provide 2.1 wt-% sulfur and thereby to simulate a hydrogen sulfide-containing atmosphere as it exists in commercial first stage hydrocracking reactors.
• sufficient cyclohexylamine was added to the feed to provide 780 wt-ppm nitrogen and thereby to simulate an ammonia-containing atmosphere as it exists in commercial first stage hydrocracking reactors.
• the selectivity values for each catalyst were total distillate (i.e., 149°C (30O 0 F) to 371°C (700 0 F)), light distillate (i.e., 149°C (300 0 F) to 288°C (550 0 F)), and heavy distillate (i.e., 288°C (550 0 F) to 371 0 C (700 0 F)).
• the relative activity value for each catalyst is entered as the difference between the required temperature of the catalyst to maintain the 65 wt-% net conversion and a reference temperature that was the same for all nine catalysts. The lower the value for relative activity, the more active is the catalyst.
• the relative catalyst activity is expressed in terms of reactor temperature above the reference temperature required to achieve 65 wt-% net conversion of the VGO to the total distillate cut.
• Catalyst A shows more total distillate selectivity, is more active, and shows a significantly higher selectivity of heavy distillate relative to light distillate compared to Catalyst B.

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