KR20130001105A - Integrated hydrocracking and dewaxing of hydrocarbons - Google Patents

Abstract

According to the present invention, an integrated process is provided for producing naphtha fuel, diesel fuel and / or lube base oil from a feedstock under sour conditions. The ability to process feedstocks under higher sulfur and / or nitrogen conditions reduces processing costs and increases flexibility in selecting an appropriate feedstock. The sour feed may be delivered to the contact dewaxing step without any separation of contaminants of sulfur and nitrogen, or may be partly removed using high pressure separation. Integrated processes include initial hydrotreatment, hydrocracking, catalytic dewaxing of the hydrolyzed effluent and any final hydrotreatment.

Description

Integrated hydrocracking and dewaxing methods of hydrocarbons {INTEGRATED HYDROCRACKING AND DEWAXING OF HYDROCARBONS}

The present invention relates to catalysts for processing sulfur and / or nitrogen containing feedstocks and to methods of using the catalysts to produce naphtha fuels, diesel fuels and lubricating oil basestocks.

Hydrocracking of hydrocarbon feedstocks is often used to convert low value hydrocarbon fractions to higher value products, such as converting vacuum gas oil (VGO) feedstocks to diesel fuels and lubricating oils. Typical hydrocracking reaction configurations may include an initial hydrotreating step, a hydrocracking step and a post-hydrogenation step. After these steps, the effluent can be fractionated and separated into the desired diesel fuel and / or lubricant basestock.

One method of fractional distillation of lubricating oil base materials is the method used by the American Petroleum Institute (API). The API Group II basestock has a saturate content of at least 90% by weight, a sulfur content of up to 0.03% by weight and a VI of greater than 80 to less than 120. The API Group III basestock is the same as the Group II basestock except that the VI is over 120. Process configurations such as those described above are typically suitable for preparing Group II and Group III basestocks from suitable feeds.

One way to improve the yield of the desired product is to use catalytic dewaxing to modify heavy molecules. Unfortunately, conventional methods of making low pour or low cloud point diesel fuel and / or lubricant basestocks suffer from different sensitivity to catalysts included in the various stages. This limits the selection of feeds that are potentially suitable for use in forming dewaxed diesel and / or Group II or higher basestocks. In conventional processing, the catalysts used for hydroprocessing and hydrolysis of oil fractions often have a relatively high tolerance to contaminants such as sulfur or nitrogen. In contrast, catalysts for catalytic dewaxing generally suffer from low tolerability to contaminants. In particular, dewaxing catalysts, which are selective for the production of high yield diesel and high yield and high VI lubricants, and which are intended primarily to operate by isomerization, are typically very sensitive to the amount of sulfur and / or nitrogen present in the feed. If contaminants are present, the activity of the dewaxing catalyst, distillate selectivity and lubricating oil yield will decrease.

In order to control the different tolerances of the catalyst, the catalytic dewaxing step is often separated from other hydroprocessing steps. In addition to the need for a separate reactor for contact dewaxing, this separation requires costly equipment and is inconvenient because it has to order the steps in the hydroprocessing sequences.

In one aspect,

Contacting the hydrotreated feedstock and hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent,

Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and

Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions

, ≪ / RTI &

The total sulfur in liquid and gaseous form supplied to the dewaxing stage is more than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,

The hydrolysis catalyst includes a zeolite Y-based catalyst,

The dewaxing catalyst comprises at least one non- dealuminated one-dimensional ten-membered ring pore zeolite, at least one Group VIII metal and at least one low surface area metal oxide refractory binder of a method for producing a naphtha fuel, diesel fuel and lubricant basestock. It is about giving.

In another aspect,

Contacting the hydrotreated feedstock and the hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent, wherein prior to said contacting, the effluent from the hydrotreating step is fed to one or more high-pressure separators to provide hydrogen Separating the gaseous fraction of the hydrotreated effluent from the liquid fraction of the treated effluent,

Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and

Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions

, ≪ / RTI &

The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,

The hydrolysis catalyst comprises a zeolite Y-based catalyst,

The dewaxing catalyst comprises at least one non- dealuminated one-dimensional ten-membered ring pore zeolite, at least one Group VIII metal and at least one low surface area metal oxide refractory binder. It is related with the provision of the manufacturing method.

In yet another aspect, the present invention,

Contacting the hydrotreated feedstock and hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent,

Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and

Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions

, ≪ / RTI &

The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,

The hydrolysis catalyst comprises a zeolite Y-based catalyst,

The dewaxing catalyst comprises at least one non- dealuminated one-dimensional ten-membered ring pore zeolite and at least one Group VIII metal, to provide a process for producing naphtha fuel, diesel fuel and lubricating oil basestocks.

In yet another aspect, the invention provides

Contacting the hydrotreated feedstock and the hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent, wherein prior to said contacting, the effluent from the hydrotreating step is fed to one or more high-pressure separators to provide hydrogen Separating the gaseous fraction of the hydrotreated effluent from the liquid fraction of the treated effluent,

Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and

Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions

, ≪ / RTI &

The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,

The hydrolysis catalyst comprises a zeolite Y-based catalyst,

A dewaxing catalyst relates to the provision of a method for producing naphtha fuel, diesel fuel and lubricating oil basestock, comprising the step of comprising at least one non- dealuminated one-dimensional 10 membered ring pore zeolite and at least one Group VIII metal.

1 is a graph of pour point of total liquid product (TLP) versus 650 ° F. + conversion.
2 is a graph of distillate yield versus 650 ° F. + conversion.
3 is a graph of naphtha yield versus 650 ° F. + conversion.
4 is a graph of lubricating oil pour point versus 700 ° F. + conversion.
Figure 5 (a) shows a prior art system for producing the dewaxed distillate / diesel fuel and lubricant basestock, Figure 5 (b) is for producing the dewaxed distillate / diesel fuel and lubricant basestock The "direct cascade" process embodiment of the present invention is shown.
FIG. 6 illustrates an embodiment of the “interstage high pressure separation” process of the present invention for preparing dewaxed distillate / diesel fuel and lubricating oil basestocks.

All numerical values within the specification and claims for carrying out the invention are modified by the indicated values "about" or "approximately," and are expected by those of ordinary skill in the art. Experimental error and variation are taken into account.

summary

In various embodiments, processes are provided for the production of lubricating oil basestocks and / or low cloud point and low pour point distillate fuels, including contact dewaxing of feeds in sour atmosphere. Sour atmosphere is an atmosphere in which the total sulfur level in liquid and gas forms exceeds 1000 ppm by weight, based on the hydrotreated feedstock. Contact dewaxing in the present invention is also referred to as hydroisomerization. The ability to perform contact dewaxing / hydroisomerization in sour atmospheres provides several advantages. The number and type of initial oil fractions available for processing can be extended because of their tolerance to contaminants at the dewaxing stage. The overall cost of the process will be lower because the ability to dewax in sour atmospheres reduces the equipment required for processing. Unlike selecting conditions that avoid exposing the dewaxing catalyst to contaminants, yields for lubricating oil and / or distillate fuel production can be improved because the processing conditions are selected to meet desired specifications. The VI of lubricant oil may also be increased. Finally, the diesel yield can be further increased by increasing the diesel end point because the pour point and / or cloud point restrictions of the diesel product are removed.

The process of the present invention involves the use of a dewaxing catalyst suitable for use in sour atmosphere with minimal conversion of high boiling point molecules to naphtha and other less valuable species. The dewaxing catalyst is used as part of an integrated process that includes initial hydrotreatment of the feed, hydrocracking the hydrotreated feed, dewaxing of the effluent from hydrocracking, and any final hydrotreatment. Since the dewaxing catalyst can withstand sour atmosphere, all of the above steps can be included in a single reactor, thus avoiding the need for costly additional reactors and other equipment to carry out this integrated process.

The dewaxing catalyst used in accordance with the present invention provides an activity advantage over conventional dewaxing catalysts in the presence of a sulfur feed. With respect to dewaxing, the sulfur feed may refer to a feed containing at least 100 ppm sulfur, or at least 1000 ppm sulfur, or at least 2000 ppm sulfur, or at least 4000 ppm sulfur, or at least 40,000 ppm sulfur. have. The feed and hydrogen gas mixture may comprise greater than 1,000 ppm by weight of sulfur, or greater than 5,000 ppm by weight of sulfur, or greater than 15,000 ppm by weight of sulfur. In yet another embodiment, sulfur may be present only in the gas, only in the liquid, or both. In the present invention, this sulfur level is defined as parts per million (ppm) of the total sulfur in liquid and gas form supplied to the dewaxing stage based on the hydrotreated feedstock.

This advantage is achieved by the use of a catalyst comprising a 10-membered ring pore one-dimensional zeolite combined with a low surface area metal oxide refractory binder, both of which are selected so that the ratio of the surface area of the micropores to the total surface area is high. Alternatively, zeolites have a low silica to alumina ratio. The dewaxing catalyst additionally comprises a metal hydrogenation function such as a Group VIII metal, preferably a Group VIII rare metal. Preferably, the dewaxing catalyst is a one-dimensional 10 membered ring pore catalyst such as ZSM-48 or ZSM-23.

The outer surface area and the surface area of the micropores are related to one way of characterizing the total surface area of the catalyst. This surface area is calculated based on the analysis of nitrogen porosity measurement data using the BET method to determine surface area (see, eg, Johnson, MFL., Jour. Catal., 52, 425 (1978)). ). The surface area of the micropores is related to the surface area by the one-dimensional pores of the zeolite in the dewaxing catalyst. Only zeolite in the catalyst will affect this part of the surface area. The external surface area may be due to the zeolite or binder in the catalyst.

Feedstock

A wide range of petroleum and chemical feedstocks can be hydroprocessed according to the present invention. Suitable feedstocks include whole and petroleum crudes, atmospheric and vacuum residues, protein deasphalted residues such as brightstock, cycle oils, FCC tower bottoms, Gas oils such as atmospheric and reduced pressure gas oils and coker gas oils, light to heavy distillates such as raw virgin distillate, hydrolysates, hydrotreated oils, dewaxed oils, slack wax features -Fischer-Tropsch wax, raffinate, and mixtures of these materials. Typical feeds include, for example, reduced pressure gas oils having a boiling point below about 593 ° C. (about 1100 ° F.) and generally in the range from about 350 ° C. to about 500 ° C. (about 660 ° F. to about 935 ° F.) The fraction of diesel fuel is therefore larger.

Feed  Initial hydrotreating

The main purpose of the hydrotreatment is typically to reduce the sulfur, nitrogen and aromatics content in the feed and is not much related to the boiling point conversion of the feed. Hydrotreating conditions include a temperature of 200 ° C. to 450 ° C. or more preferably 315 to 425 ° C., 250 to 5000 psig (1.8 MPa to 34.6 MPa) or more preferably 300 to 3000 psig (2.1 MPa to 20.8 MPa), 0.2 to 10h ^-1 liquid hourly space velocity (LHSV: liquid hourly space velocity) and 200 to ^{10,000scf / B (35.6m 3 / m} 3 to about 1781m ³ / m ³⁾ or more preferably from 500 to 10,000scf / B ( 89 m ³ / m ³ to 1781m ³ / m ³ ) of the hydrotreating ratio. Hydrotreating catalysts are typically those containing Group VIB metals (based on a periodic table published by Fisher Scientific) and non-rare metals Group VIII metals, ie iron, cobalt and nickel and mixtures thereof. . Such metals or mixtures of metals are typically present as oxides or sulfides on refractory metal oxide supports. Suitable metal oxide supports include low acidic oxides such as silica, alumina or titania, preferably alumina. Preferred aluminas have an average pore size of 50 to 200 mm 3, preferably 75 to 150 mm 3, a surface area of 100 to 300 m ² / g, preferably 150 to 250 m ² / g, a pore volume of 0.25 to 1.0 cm ³ / g, Preferably porosity alumina, such as gamma or eta, from 0.35 to 0.8 cm ³ / g. The support is preferably not promoted with a halogen such as fluorine because it generally increases the acidity of the support.

Preferred metal catalysts are cobalt / molybdenum (Co 1-10% as oxide, Mo 10-40% as oxide), nickel / molybdenum (Ni 1-10% as oxide, Mo 10-40% as oxide) or nickel on alumina Tungsten (1-10% Ni as oxide, W 10-40% as oxide). Particular preference is given to stacked beds and Nebula-20 of nickel / molybdenum catalysts such as KF-840, KF-848 or KF-848 or KF-840.

Alternatively, the hydrotreating catalyst may be a bulk metal catalyst, or a combination on a stack of supported bulk metal catalysts. Bulk metal means that the catalyst is not supported, wherein the bulk catalyst particles, when calculated as metal oxide, contain from 30 to 100 at least one Group VIII non-rare metal and at least one Group VIB metal, based on the total weight of the bulk catalyst particles. In weight percent, the bulk catalyst particles have a surface area of at least 10 m ² / g. Moreover, the bulk metal hydrotreatment catalysts used herein, when calculated as metal oxides, contain from about 50% to about 100% by weight and even more of at least one Group VIII non-rare metal and at least one Group VIB metal, based on the total weight of the particles. Preferably from about 70 to about 100% by weight. The amounts of Group VIB and Group VIII non-rare metals can be easily determined by VIB TEM-EDX.

Preference is given to bulk catalyst compositions comprising one Group VIII non-rare metal and two Group VIB metals. In this case, the bulk catalyst particles were found to be sinter-resistant. Thus, the active surface area of the bulk catalyst particles is maintained during use. The molar ratio of Group VIB to VIII non-rare metals generally ranges from 10: 1 to 1:10 and preferably from 3: 1 to 1: 3. In the case of core-shell structured particles, the ratio, of course, applies to the metal contained in the shell. If more than one Group VIB metal is contained in the bulk catalyst particles, the ratio of different Group VIB metals is generally not of great importance. The same holds true if more than one Group VIII non-rare metal is used. When molybdenum and tungsten are present as the Group VIB metals, the molybdenum: tungsten ratio is preferably 9: 1 to 1: 9. Preferably, the Group VIII non-rare metals comprise nickel and / or cobalt. It is also preferred that the Group VIB metal comprises a combination of molybdenum and tungsten. Preferably, a combination of nickel / molybdenum / tungsten and cobalt / molybdenum / tungsten and nickel / cobalt / molybdenum / tungsten is used. Precipitates of this type appear to be sinter-resistant. Thus, the active surface area of the precipitate is maintained during use. Preferably, the metal is present as an oxidizable compound of the corresponding metal or, if the catalyst composition is sulfided, as a sulfidable compound of the corresponding metal.

In addition, the bulk metal hydrotreating catalyst used herein preferably has a surface area of at least 50 m ² / g and more preferably at least 100 m ² / g. In addition, the pore size distribution of the bulk metal hydrotreating catalyst is preferably about the same as the conventional hydrotreating catalyst. More particularly, such bulk metal hydrotreatment catalysts have a pore volume of 0.05 to 5 ml / g, more preferably 0.1 to 4 ml / g, still more preferably 0.1, as measured by hydrogen adsorption. To 3 ml / g and most preferably 0.1 to 2 ml / g. Preferably, no pores smaller than 1 nm are present. Moreover, such bulk metal hydrotreating catalysts preferably have a median diameter of at least 50 nm, more preferably at least 100 nm, and preferably at most 5000 μm and even more preferably at most 3000 μm. Even more preferably, the median diameter is in the range from 0.1 to 50 μm, most preferably in the range from 0.5 to 50 μm.

Hydrolysis  fair

Hydrocracking catalysts typically contain an acid support such as amorphous silica, alumina, a zeolite for decomposition such as USY, a nonmetal sulfide on an acidified alumina. Often these acidic supports are mixed or combined with other metal oxides such as alumina, titania or silica.

The hydrocracking process includes a temperature of about 200 ° C. to about 450 ° C., a hydrogen pressure of about 250 psig to about 5000 psig (1.8 MPa to 34.6 MPa), a liquid hourly space velocity of about 0.2 h ⁻¹ to about 10 h ⁻¹ , and about 35.6 m ³ / m ³ to about 1781m ³ / m ³ (about 200SCF / B to about 10,000SCF / B) may be carried out at a hydrogen treatment gas ratio. Typically, in most cases, the conditions of hydrogen pressure, about 0.5h ^-1 to about 2h ^-1 of about 300 ℃ to about 450 ℃ temperature range of about 500psig to about 2000psig (3.5MPa to about 13.9MPa) a liquid hourly Space velocity and hydrotreating gas ratio of about 213m ³ / m ³ to about 1068m ³ / m ³ (about 1200SCF / B to about 6000SCF / B).

Taro  fair

The product from hydrocracking is then cascaded directly into the catalytic dewaxing reaction zone. Unlike conventional processes, no separation is required between hydrocracking and the contact dewaxing stage. Omission of the separation step provides various effects. With regard to the separation itself, no additional equipment is required. In one form, the contact dewaxing stage and the hydrolysis stage are located in the same reactor. Alternatively, the hydrocracking and catalytic dewaxing processes can be carried out in separate reactors. By eliminating the separation step there is also no need to recompress the feed. Instead, the effluent from the hydrocracking stage can be maintained at the processing pressure as the effluent is sent to the dewaxing stage.

Omitting the separation step between hydrocracking and catalytic dewaxing also means that any sulfur in the feed of the hydrocracking stage will still be present in the effluent passing from the hydrocracking stage to the catalytic dewaxing stage. Some of the organic sulfur in the feed to the hydrocracking stage will be converted to H ₂ S during hydrotreating. Similarly, organic nitrogen in the feed will be converted to ammonia. However, without the separation step, H ₂ S and NH ₃ formed during the hydrotreatment will move to the contact dewaxing stage with the effluent. In addition, the absence of the separation step means that any light gas (C ₁ -C ₄ ) formed during hydrocracking will still be present in the effluent. The total sulfur from the hydrotreating process in both the organic liquid form and the gas phase (hydrogen sulfide) is greater than 1,000 ppm, or at least 2,000 ppm, or at least 5,000 ppm, or at least 10,000 ppm, or at least 20,000 ppm, or 40,000 It may be at least ppm by weight. In the present invention, this sulfur level is defined in parts per million parts by weight of the total sulfur in liquid and gaseous form fed to the dewaxing stage, based on the hydrotreated feedstock.

The omission of the separation step between hydrocracking and catalytic dewaxing may be partially possible by the ability of the dewaxing catalyst to maintain catalytic activity in the presence of elevated levels of nitrogen and sulfur. Conventional catalysts often require pretreatment of the feedstream to reduce the sulfur content to less than several hundred ppm. In contrast, hydrocarbon feedstreams containing up to 4.0 wt.% Sulfur can be effectively processed using the catalyst of the present invention. In one embodiment, the total sulfur content in the liquid and gaseous form of the hydrogen containing gas and the hydrotreated feedstock is at least 0.1%, or at least 0.2%, or at least 0.4%, or at least 0.5%, or 1% Or at least 2 weight percent, or at least 4 weight percent. Sulfur content can be measured by standard ASTM method D2622.

In an alternative embodiment, a simple flash high pressure separation step without stripping can be carried out without depressurizing the feed in the effluent from the hydrotreating reactor. In this embodiment, the high pressure separation step removes any gaseous sulfur and / or nitrogen contaminants from the gaseous effluent. However, since the separation takes place at a pressure comparable to the process pressure in the hydrotreating or hydrocracking step, the effluent will still contain a substantial amount of dissolved sulfur. For example, the amount of dissolved sulfur in H ₂ S form may be at least 100 vppm, or at least 500 vppm, or at least 1000 vppm, or at least 2000 vppm, or at least 5000 vppm, or at least 7000 vppm.

The hydrotreating gas circulation loop and the constituent gases can be arranged and regulated in various ways. In the direct cascade, the process gas can be fed to a hydrotreating reactor and flowed from a high pressure flash drum at the rear end of the hydrocracking and / or dewaxing section of the apparatus or circulated by a compressor. In a simple flash arrangement, the process gas may be fed in parallel in the hydrotreating and hydrolysis and / or dewaxing reactors in both perfusion or circulating modes. In a circulating manner, the constituent gases can be injected into the apparatus anywhere in the high pressure circuit, preferably into the hydrocracking / dropping reactor zone. In a circulating manner, the process gas may be scrubbed with amine or any other suitable solution to remove H ₂ S and NH ₃ . In another form, the process gas can be recycled without washing or scrubbing. Alternatively, the liquid effluent may be combined with any hydrogen containing gas including but not limited to H ₂ S containing gas.

Preferably, the dewaxing catalyst according to the invention is a zeolite, which basically performs dewaxing by isomerizing the hydrocarbon feedstock. More preferably, the catalyst is a zeolite having a one-dimensional pore structure. Suitable catalysts include 10 membered ring pore zeolites such as EU-1, ZSM-35 (or ferrite), ZSM-11, ZSM-57, NU-87, SAPO-11 and ZSM-22. Preferred materials are EU-2, EU-11, ZBM-30, ZSM-48 or ZSM-23. ZSM-48 is most preferred. It should be noted that zeolites having a ZSM-23 structure having a silica to alumina ratio of about 20: 1 to about 40: 1 may sometimes be referred to as SSZ-32. Other molecular sieves of the isoform structure in this material include theta-1, NU-10, EU-13, KZ-1 and NU-23.

In various embodiments, the catalyst according to the present invention further comprises a metal hydrogenation component. The metal hydrogenation component is typically a Group VI and / or Group VIII metal. Preferably, the metal hydrogenation component is a Group VIII rare metal. More preferably, the metal hydrogenation component is Pt, Pd or mixtures thereof.

The metal hydrogenation component can be added to the catalyst in any convenient manner. The technique of adding metal hydrogenation components is due to incipient wetness. For example, after combining the zeolite and the binder, the combined zeolite and binder can be extruded into catalyst particles. These catalyst particles can then be exposed to a solution containing the appropriate metal precursor. Alternatively, metal can be added to the catalyst by ion exchange, where the metal precursor is added to the zeolite (or zeolite and binder) mixture prior to extrusion.

The amount of metal in the catalyst may be at least 0.1%, or at least 0.15%, or at least 0.2%, or at least 0.25%, or at least 0.3%, or at least 0.5% by weight, based on the catalyst. The amount of metal in the catalyst can be up to 5 wt%, or up to 2.5 wt%, or up to 1 wt%, or up to 0.75 wt%, based on the catalyst. In embodiments where the metal is Pt, Pd, another Group VIII rare metal, or a combination thereof, the amount of metal is preferably 0.1 to 2% by weight, more preferably 0.25 to 1.8% by weight, and even more preferably 0.4 To 1.5% by weight.

Preferably, the dewaxing catalyst used in the process according to the invention is a catalyst with a low ratio of silica to alumina. For example, in ZSM-48, the ratio of silica to alumina in the zeolite can be less than 200: 1, or less than 110: 1, or less than 100: 1, or less than 90: 1, or less than 80: 1. In a preferred embodiment, the ratio of silica to alumina can be 30: 1 to 200: 1, 60: 1 to 110: 1, or 70: 1 to 100: 1.

Wax dewaxes useful in the process according to the invention may also comprise a binder. In some embodiments, the dewaxing catalyst used in the process according to the invention is formulated using a low surface area binder, the low surface area binder having a surface area of 100 m ² / g or less, or 80m ² / g or less, or 70m ² / g or less Means.

Alternatively, the binder and zeolite particle sizes are chosen to provide a catalyst having a desired ratio of the surface area of the micropores to the total surface area. In the dewaxing catalyst used according to the invention, the surface area of the micropores corresponds to the surface area from the one-dimensional pores of the zeolite in the dewaxing catalyst. The total surface area corresponds to the sum of the surface area of the micropores and the external surface area. Any binder used in the catalyst will not affect the surface area of the micropores and will not significantly increase the total surface area of the catalyst. The external surface area represents the surface area of the total catalyst minus the surface area of the micropores. Both binders and zeolites can affect the value of the external surface area. Preferably, the ratio of the surface area of the micropores to the total surface area in the dewaxing catalyst will be at least 25%.

Zeolites may be combined with the binder in any convenient manner. For example, the combined catalyst starts with a powder of both zeolite and binder, and combines the powder with added water to mix well to form a mixture, which is then extruded to produce a combined catalyst of the desired size. Can be formed. Extrusion aids may also be used to modify the extrusion flow characteristics of the zeolite and binder mixtures. The amount of skeletal alumina in the catalyst can range from 0.1 to 3.33 wt%, or from 0.2 to 2 wt%, or from 0.3 to 1 wt%.

In yet another embodiment, binders composed of two or more metal oxides may also be used. In this embodiment, the weight percent of the low surface area binder is preferably greater than the weight percent of the high surface area binder.

Alternatively, if the two metal oxides used to form the mixed metal oxide binder have sufficiently low surface areas, the fraction of each of the metal oxides in the binder is less important. When two or more metal oxides are used to form the binder, the two metal oxides may be incorporated into the catalyst by any convenient method. For example, one binder may be mixed with the zeolite during the formation of the zeolite powder, for example during spray drying. The spray dried zeolite / binder powder may then be mixed with the second metal oxide binder before being extruded.

In yet another embodiment, the dewaxing catalyst is self-bonding and contains no binder.

Process conditions in the contact dewaxing zone are at a temperature of 200 ° C. to 450 ° C., preferably 270 to 400 ° C., a hydrogen partial pressure of 1.8 to 34.6 mPa (250 to 5000 psig), preferably 4.8 to 20.8 mPa, and 0.2 to 10 v / v. / hr, preferably from 0.5 to 3.0v / v / hr of a liquid hourly space velocity, and 35.6 to about 1781m ³ / m ³ (200 to 10,000scf / B), preferably from 178 to 890.6m ³ / m ³ (1000 to Hydrogen cycle rate of 5000 scf / B).

Post-hydrogenation

The effluent from the dewaxing stage can then optionally be treated in the final hydrotreating step. The catalyst during the hydrotreating step may be the same as described above for the first hydrotreatment. The reaction conditions for the second hydrotreating step may also be similar to the conditions for the first hydrotreating.

After post-hydrogenation, the various fractions of the effluent may be suitable for use as diesel fuel or lubricant basestock. However, in some embodiments, the lubricant basestock formed may only be partially dewaxed. In this embodiment, further processing may be necessary for the fraction desired for use as a lubricant basestock. For example, after the post-hydrogenation step, the effluent may be fractionally distilled to yield a diesel fuel fraction and a lubricating oil basestock fraction. The lubricating oil basestock fraction can then be treated in a solvent dewaxing step or another contact dewaxing step to obtain the desired properties of the lubricating oil basestock. The lubricating oil basestock fraction can then be hydrofinishing and reduced pressure stripping.

fair Example  One

In one embodiment, the effluent from the hydrotreatment step can be cascaded directly into the hydrolysis step. Hydrotreating and hydrocracking catalysts may be located in a single reactor. This may be referred to herein directly as the cascade embodiment (see FIG. 5B). Depending on the choice of other catalysts and reaction conditions, the products of the process may exhibit improved viscosity, viscosity index, saturate content, low temperature properties, volatility and depolarity. The reactor can also be operated in any suitable catalytic bed arrangement, for example fixed bed, slurry bed or ebulating bed, but fixed bed, cocurrent downflow is generally used. In embodiments in which the effluent from the hydrotreatment step is cascaded directly into the hydrolysis step, the conditions in the hydrotreatment step may be selected to match the conditions in the hydrolysis step.

FIG. 5 schematically shows a comparison between a conventional reaction system (FIG. 5A) and one reaction system suitable for the practice of the present invention (FIG. 5B). 5A shows a prior art reaction system having a conventional reactor for carrying out a hydrolysis reaction.

5B illustrates one embodiment of the reaction system of the present invention for performing a direct cascade process. The initial phase of the reactor includes a hydrotreating catalyst for removing heteroatomic contaminants from the feed. The feed is then exposed to the hydrocracking catalyst, preferably without intermediate separation. After hydrocracking, the effluent from the hydrocracking stage is exposed to the dewaxing catalyst without intermediate separation. After dewaxing, the effluent from the dewaxing step is exposed to the second hydrotreating catalyst and the undesirable saturate olefinic species for further removal of heteroatoms.

In a conventional prior art schematic, the catalytic contact dewaxing and / or catalyst isomerization is carried out in a separate reactor. This is because conventional catalysts are contaminated with heteroatom contaminants (eg, H ₂ S, NH ₃ , organic sulfur and / or organic nitrogen) that are typically present in hydrocracker effluents. Thus, in the conventional schematics, a separation step is used to first reduce the amount of heteroatomic contaminants. In addition, since distillation needs to be carried out to separate the various cuts from the hydrocracker effluent, the distillation can be carried out simultaneously with the distillation and thus before the dewaxing. This means that some valuable hydrocarbon molecules that could be used in diesel or lubricating oil fractions are lost.

In the direct cascade embodiment of FIG. 5B, a layer of dewaxing catalyst is included between the hydrolysis step and the final hydrotreatment. By using a contaminant tolerant catalyst, a moderate dewaxing step can be performed on the entire effluent from the hydrocracking step. This means that all molecules present in the hydrocracking effluent are exposed to moderate dewaxing. Such mild dewaxing will change the boiling point of the long chain molecules, thus enabling the molecules that are normally released in the distillation step as bottoms to be converted into molecules suitable for the lubricant basestock. Similarly, some molecules suitable for the lubricant basestock will be converted to diesel range molecules. The effective effect is that, unlike separation into bottoms that are likely to decompose in gasoline, more hydrocracker effluents will be incorporated into high value products. The amount of diesel and / or lubricant basestock should also increase with the nature of the feedstock.

In FIG. 5B, the first hydrotreatment step, the hydrolysis step, the sour service dewaxing step and the second hydrotreatment step are carried out in the same reactor. There is an advantage of minimizing the number of reactors. Alternatively, each of these steps can be carried out in a separate reactor. For example, the hydrocracking step can be carried out in one reactor and the subsequent sour service dewaxing step can be carried out in a separate reactor without any separation between the two reactors.

fair Example  2

In an alternative embodiment, the effluent from the hydrotreating step can be passed through a high pressure separator to remove H ₂ S and NH ₃ before the subsequent hydrolysis step. This may be referred to herein as an “interstage high pressure separation” embodiment (see FIG. 6). The interstage high pressure separation mode can result in higher conversion in the downflow hydrocracking / hydrogenation reactor. Figure 6 schematically illustrates one embodiment of a reaction system of the present invention for carrying out an interstage high pressure separation process. 6 schematically shows an arrangement of hydrotreating reactor 720 and subsequent high pressure separation apparatus. In FIG. 6, the total effluent from hydrotreating reactor 720 passes to one or more high pressure separation devices, such as high pressure separator 722 and 723 pairs. The high pressure separation device separates the gaseous fraction of the effluent from the liquid phase fraction. The resulting effluent 734, which contains dissolved H ₂ S and possibly organic sulfur, is then remixed with the hydrogen containing gas. The hydrogen containing gas may contain H ₂ S. The combined mixture is then sent to another reactor containing a hydrocracking catalyst. After hydrocracking, the effluent from the hydrocracking stage is exposed to the sour service dewaxing catalyst for isomerization without intermediate separation. In one form, the hydrocracking catalyst and the dewaxing catalyst are located in the same reactor. The effluent from the dewaxing stage can then be treated, optionally in the final hydrotreating step, and then separated into several fractions by means of a fractionator. Such fractions may include, for example, light fuel type products such as naphtha fractions, light fuel type products such as diesel fractions and heavy lubricating oil basestock fractions. The lubricating oil basestock fraction may then be treated in a solvent dewaxing step or another contact dewaxing step to obtain the desired properties for the lubricating oil basestock. The lubricating oil basestock fraction may then be hydrofinishing and stripped under reduced pressure. High pressure separation will remove some gaseous sulfur and nitrogen from the effluent, which is removed as sour gas stream 732 for further processing. However, the separated effluent 734 passing through the dewaxing stage may still contain greater than 1000 ppm by weight of total sulfur in liquid and gaseous form, for example based on the hydrotreated feedstock. Since the dewaxing catalyst is exposed to less severe sour atmosphere, this partial reduction in sulfur and nitrogen content in the effluent may improve the activity and / or shelf life of the dewaxing catalyst.

In another form, the hydrocracking catalyst and the dewaxing catalyst are located in two separate reactors without intermediate separation. After the sour service dewaxing catalyst, the dewaxed hydrocracking effluent may be transported to the second hydrotreating catalyst and the undesirable saturate olefinic species for further removal of heteroatoms. The second hydrotreating step may be located in the same reactor as the hydrocracking and dewaxing step or may be located in a separate downflow reactor. After the final hydrotreatment step, the effluent is separated into several fractions by means of a fractionator. Such fractions may include, for example, light fuel type products such as naphtha fractions, light fuel type products such as diesel fractions and heavy lubricating oil basestock fractions. The lubricant basestock fraction is then subjected to a solvent dewaxing step or another contact dewaxing step in order to obtain the desired properties for the lubricant basestock. The lubricating oil basestock fraction may then be hydrofinishing and stripped under reduced pressure.

Taro  Catalytic synthesis

In one form of the invention, the catalytic dewaxing catalyst comprises from 0.1% to 3.33% by weight of skeletal alumina, from 0.1% to 5% by weight of Pt, the SiO ₂ : Al ₂ O ₃ ratio from 200: 1 to 30: 1, and At least one low surface area refractory metal oxide binder having a surface area of 100 m ² / g or less.

One example of a molecular sieve suitable for use in the claimed invention is ZSM-48 having an SiO ₂ : Al ₂ O ₃ ratio of less than 110, preferably about 70 to about 110. In the following embodiments, the ZSM-48 crystals are "as-synthesized" crystals that still contain an organic template (SiO ₂ : Al ₂ O ₃ ratio of 200: 1 or less): calcined Crystals such as Na-type ZSM-48 crystals; Or calcined and ion-exchanged crystals, such as H-type ZSM-48 crystals.

The ZSM-48 crystals after removal of the structural directing agent have a specific morphology and molar composition according to the following structural formula:

(n) SiO ₂ : Al ₂ O ₃

In the above structural formula, n is 70 to 110, preferably 80 to 100, more preferably 85 to 95. In one embodiment n is at least 70, or at least 80, or at least 85. In yet another embodiment, n is 110 or less, or 100 or less, or 95 or less. In still other embodiments, Si may be substituted with Ge and Al may be substituted with Ga, B, Fe, Ti, V and Zr.

ZSM-48 crystals in the form of synthesized are prepared from a mixture with silica, alumina, base and hexamethonium salt directing agents. In one embodiment, the molar ratio of structural directing agent: silica in the mixture is less than 0.05, or less than 0.025, or less than 0.022. In another embodiment, the molar ratio of structural directing agent: silica in the mixture is at least 0.01, or at least 0.015, or at least 0.016. In yet another embodiment, the molar ratio of structural directing agent: silica in the mixture is 0.015 to 0.025, preferably 0.016 to 0.022. In one embodiment, the ZSM-48 crystals in the form of synthesized have a molar ratio of silica: alumina between 70 and 110. In yet another embodiment, the ZSM-48 crystals in the form of synthesized have a molar ratio of silica: alumina of at least 70, or at least 80, or at least 85. In yet another embodiment, the ZSM-48 crystals in the form of synthesized have a molar ratio of silica: alumina of 110 or less, or 100 or less, or 95 or less. For certain preparations of ZSM-48 crystals in the form of synthesized bars, the molar composition will include silica, alumina and fragrances. It should be noted that the ZSM-48 crystals in the form of synthesized bars may have a molar ratio slightly different from the molar ratio of the reactants in the reaction mixture used to prepare the synthesized bar. This result can occur because 100% of the reactants in the reaction mixture do not fully incorporate into the formed crystals (from the reaction mixture).

ZSM-48 compositions are prepared from aqueous reaction mixtures comprising salts of silica or silicates, salts of alumina or soluble aluminates, bases and directing agents. To achieve the desired crystal morphology, the reactants in the reaction mixture have the following molar ratios:

SiO ₂ : Al ₂ O ₃ (preferred) = 70 to 110

H ₂ O: SiO ₂ = 1 to 500

OH ^-: SiO ₂ = 0.1 to 0.3

OH ^-: SiO ₂ (preferably also) = 0.14 to 0.18

Template: SiO ₂ = 0.01 to 0.05

Template: SiO ₂ (preferred) = 0.015 to 0.025

In this ratio, two ranges are provided for both the ratio of base: silica and the ratio of structure directing agent: silica. At this ratio, a broader range includes mixtures that result in the formation of ZSM-48 crystals with a predetermined amount of Kenyaite and / or acicular morphology. Preferred ranges should be used in situations where kenyanite and / or acicular morphology is not desired.

The silica source is preferably precipitated silica and is available from Degussa. Other silica sources include silica powders such as precipitated silicas such as Zeosil ^® and silica gels, silicate colloidal silicas such as Ludox ^® or dissolved silicas. In the presence of a base, these other silica sources can form silicates. Alumina is in the form of a soluble salt, preferably sodium salt, and is available from US Aluminate. Other suitable aluminum sources include other aluminum salts, such as chlorides thereof, aluminum alcoholates, or hydrated aluminas such as gamma alumina, pseudobohemite, and colloidal alumina. The base used to dissolve the metal oxide can be any alkali metal hydroxide, preferably sodium or potassium hydroxide, ammonium hydroxide, di-quad hydroxide, and the like. The directing agent is a hexamethonium salt such as hexamethonium dichloride or hexamethonium hydroxide. Anions (in addition to chlorides) may be other anions such as hydroxides, nitrates, sulfates, other halides and the like. Hexamethonium dichloride is N, N, N, N ', N', N'-hexamethyl-1,6-hexanediammonium dichloride.

In one embodiment, the crystals obtained from the synthesis according to the invention have a morphology free of fibrous morphology. Fibrous morphology is undesirable because its crystalline morphology inhibits the contact dewaxing activity of ZSM-48. In another embodiment, the crystals obtained from the synthesis according to the invention have morphologies containing a low percentage of acicular morphology. The amount of acicular morphology present in the ZSM-48 crystals can be up to 10%, or up to 5%, or up to 1%. In an alternative embodiment, the ZSM-48 crystals may be free of acicular morphology. Since needle crystals are believed to reduce the activity of ZSM-48 for some types of reactions, it is desirable for some applications to have low content of needle crystals. In order to obtain the high purity of the desired morphology, the ratio of silica: alumina, base: silica and fragrance: silica in the reaction mixture should be used in accordance with embodiments of the present invention. In addition, if a composition free of Kenyanite and / or needle-free morphology is desired, the preferred range should be used.

The ZSM-48 crystals as synthesized should be at least partially dried before use or before further treatment. Drying can be carried out by heating at a temperature of 100 to 400 ° C, preferably 100 to 250 ° C. The pressure can be atmospheric or subatmospheric. When drying is carried out under partially reduced pressure conditions, the temperature may be lower than the temperature at atmospheric pressure.

Typically, the catalyst is combined with a binder or matrix material prior to use. The binder is frictional and resistant to the desired temperature of use. The binder may be catalytically active or inert and includes other zeolites, other inorganic materials such as clays and metal oxides such as alumina, silica, titania, zirconia and silica-alumina. Clay is kaolin, bentonite and montmorillonite, and commercially available ones can be used. These may be blended with other materials, such as silicates. In addition to silica-alumina, other porous matrix materials include biphasic materials such as silica-magnesia, silica-toria, silica-zirconia, silica-berilia and silica-titania, as well as three-phase materials such as silica-alumina-magnesia, Silica-alumina-toria and silica-alumina-zirconia. The matrix may be in the form of a co-gel. The bound ZSM-48 backbone alumina will range from 0.1% to 3.33% backbone alumina.

As part of the catalyst, ZSM-48 crystals can also be used with the metal hydrogenation component. The metal hydrogenation component may be selected from Groups 6-12, preferably Groups 6 and 8-10 of the Periodic Table based on the IUPAC system having Groups 1-18. Examples of such metals include Ni, Mo, Co, W, Mn, Cu, Zn, Ru, Pt or Pd, preferably Pt or Pd. Mixtures of metal hydrides such as Co / Mo, Ni / Mo, Ni / W and Pt / Pd, preferably Pt / Pd, may also be used. The amount of metal hydride or metals may range from 0.1 to 5% by weight, based on the catalyst. In one embodiment, the amount of metal or metals is at least 0.1 weight percent, or at least 0.25 weight percent, or at least 0.5 weight percent, or at least 0.6 weight percent, or at least 0.75 weight percent, or at least 0.9 weight percent. In another embodiment, the amount of metal or metals is at most 5%, or at most 4%, or at most 3%, or at most 2%, or at most 1%. Methods of filling metal onto ZSM-48 catalysts are well known and include, for example, impregnating and heating the ZSM-48 catalyst with a metal salt of the hydrogenation component. ZSM-48 catalysts containing metal hydrides may also be sulfided prior to use.

The high purity ZSM-48 crystals prepared according to this embodiment have a relatively low ratio of silica: alumina. This low silica: alumina ratio means that the catalyst is more acidic. Despite this increased acidity, the catalysts have excellent yields as well as superior activity and selectivity. The catalyst also provides environmental benefits in terms of the effect on health from the crystalline form, and small crystal sizes are also advantageous for catalytic activity.

In the catalyst according to the invention incorporating ZSM-23, any suitable method for producing ZSM-23 having a low SiO ₂ : Al ₂ O ₃ ratio may be used. U. S. Patent No. 5,332, 566 provides an example of a synthesis method suitable for preparing ZSM-23 having a low SiO ₂ : Al ₂ O ₃ ratio. For example, the methylation of iminobispropylamine with an excess of iodomethane can form a suitable directing agent for preparing ZSM-23. Methylation is achieved by dropwise addition of iodomethane to iminobispropylamine solvated in absolute ethanol. The mixture is heated at reflux temperature of 77 ° C. for 18 hours. The resulting solid product is filtered off and washed with absolute ethanol.

The fragrance prepared by the method can then be mixed with a colloidal silica sol (30% SiO ₂ ), an alumina source, an alkali cation source (eg Na or K) and deionized water to form a hydrogel. The alumina source can be any convenient source, such as alumina sulfate or sodium aluminate. The solution is then heated to a crystallization temperature such as 170 ° C. and the formed ZSM-23 crystals are dried. The ZSM-23 crystals can then be combined with a low surface area binder to form the catalyst according to the invention.

Examples of the present invention are presented below, which should not be construed as limiting.

Example

Example  1A: SiO ₂ Of Al ₂ O ₃  Ratio is about 70/1 Morphology  Have ZSM Synthesis of -48 Crystals

A mixture of deionized (DI) water, hexamethonium chloride (56% solution), ultrasil silica, sodium aluminate solution (45%) and 50% sodium hydroxide solution and about 0.15% (relative to the reaction mixture) The mixture was prepared from ZSM-48 seed crystals. The mixture has the following molar composition:

The mixture was reacted in a 5-gal autoclave at 320 ° F. (160 ° C.) with stirring at 250 RPM for 48 hours. The product was filtered off, washed with deionized (DI) water and dried at 250 ° F. (120 ° C.). The XRD pattern of the synthesized material showed a typical pure phase of ZSM-48 topology. SEM of the material as synthesized shows that the material consists of small irregular shaped crystal masses (average crystal size about 0.05 microns). The formed ZSM-48 crystals were SiO ₂ / Al ₂ O ₃ The molar ratio was about 71. The synthesized crystals were converted to hydrogen form by ion exchange with ammonium nitrate solution three times at room temperature, then dried at 250 ° F. (120 ° C.) and calcined at 1000 ° F. (54 ° C.) for 4 hours. The ZSM-48 (70: 1 SiO ₂ : Al ₂ O ₃ ) crystals formed had a total surface area of about 290 m ² / g (outer surface area of about 130 m ² / g) and an alpha value of about 100, resulting in current ZSM-48 (90 : 1 SiO ₂ : Al ₂ O ₃ ) was about 40% higher than alumina crystals. The H-form crystals were then steamed at 700 ° F., 750 ° F., 800 ° F., 900 ° F. and 1000 ° F. for 4 hours to enhance activity, and the alpha values of these treated products were as follows:

170 (700 ° F), 150 (750 ° F), 140 (800 ° F), 97 (900 ° F) and 25 (1000 ° F).

Example  1B: sour  service Taro  Preparation of the catalyst

65 wt% ZSM-48 (about 70/1 SiO ₂ / Al ₂ O ₃ , see Example 1A) was mixed with 35 wt% P25 TiO ₂ binder and extruded into 1/20 ”tetralobes sour service hydroisomerization catalyst The catalyst was then precalcined in nitrogen at 1000 ° F., ammonium exchange with ammonium nitrate, and calcined in full air at 1000 ° F. The extrudate was then full steam at 750 ° F. steam for 3 hours The initial impregnation with platinum tetraamine nitrate impregnated the 0.6 wt% platinum steamed catalyst, dried and then calcined at 680 ° F. for 3 hours in air. The ratio of the surface area of the micropores to the total surface area is about 45%.

Example  2: sour  service Hydrolysis Process evaluation of hydrogen isomerization

In this example, the benefits of replacing some of the hydrocracking (HDC) catalysts with sour service hydroisomerization (HI) catalysts are evaluated. The hydrocracking catalyst used in this study was a zeolite Z-3723 catalyst.

As shown in Table 1, the reactors, i.e., two reactors in series, were charged to replace about 50% of the hydrocracking (HDC) catalyst with the sour service hydroisomerization (HI) catalyst described in Example 1. Evaluated. Feed characteristics of the MVGO are described in Table 2 below.

Reactor Fill Composition Reactor ＃ 1 basic
HDT / HDC / HDT Sour service
HDT / HDC / HI / HDT  -KF-848 (hydrogen treatment) 40% 40% Reactor ＃ 2  -KF-848 (hydrogen treatment) 30% 30%  Zeolite Z-3723 (hydrogen decomposition) 25% 12.5% 0.6 wt% Pt (hydroisomerization) on 35/65 TiO ₂ / ZSM-48 - 12.5%  -KF-848 (hydrogen treatment) 5% 5%

MVGO Feed Characteristics Feed characteristics MVGO Feed 700 ° F + in feed 90 Feed Pour Point, ℃ 30 Solvent dewaxed oil feed pour point, ℃ -19 Solvent dewaxed oil feed 100 ° C viscosity, cSt 7.55 Solvent Wax Oil Feed VI 57.8 Organic Sulfur in Weight Feed (ppm) 25,800 Organic Nitrogen in Feed (ppm) 809

The catalyst was first heated at 25 ° F./hr at 800 psig and dried in hydrogen at 225 ° F. Once the reaction temperature reached 225 ° F., the spiked feed (DMDS mixed with LGO at 2.3 wt.% S) was introduced at 800 psig at a ratio of 1 LHSV and 1000 scf / B hydrogen gas to feed. After the catalyst was infiltrated for 3 hours, the reactor was heated from 40 ° F / hr to 450 ° F. The temperature was then maintained at 450 ° F. for about 10 hours. A second spiked feed (DMDS mixed with MVGO at 2.5% by weight S) was introduced at 800 psig and 450 ° F. at a ratio of 1 LHSV and 1000 scf / B hydrogen gas to feed. After 1 hour, the reactor temperature was increased from 40 ° F./hr to 610 ° F. The temperature was kept at 610 ° F. for about 5 hours. The reactor was then heated to 664 ° F. at 40 ° F./hr and held at 664 ° F. for 15 hours. After 15 hours, sulfidation was complete, the MVGO feed was introduced into the apparatus, and conditions were adjusted to achieve a conversion of about 40%. During the evaluation, reactor # 2 was operated at a temperature about 25 ° F. higher than reactor # 1 to simulate a commercial temperature profile. The sour service hydroisomerization catalyst was not specifically dried or pre-reduced before being charged to the reactor and went through the same activation procedure as the hydrotreating and hydrocracking catalysts described above.

Process conditions, conversions, yields and total liquid product properties are summarized in Table 3. In the basic case only the hydrocracking catalyst is included, whereas in the HDC / HI case the hydrolysis catalyst and the hydroisomerization catalyst are included in a single reactor.

Pilot device evaluation Condition 1 Condition 2 Condition 3 Condition 4 basic HDC / HI basic HDC / HI basic HDC / HI HDC / HI Equal Supply Ratio, KBD 35 35 35 35 Temperature of reactor ＃ 1, ℉ 680 690 705 720 Temperature of reactor ＃ 2, ℉ 705 715 730 740 Treatment Gas Rate, SCF / B To 4000 To 4000 To 4000 To 4000 Full LHSV, 1 / hr 0.75 0.75 0.75 0.75 Pressure, psig To 1250 To 1250 To 1250 To 1250 650 ° F +, conversion rate, weight% 26 14 26.5 17.0 39.7 29.0 43 Yield, Volume% Gas (C _4- ), weight% 0.4 0.6 0.9 1.0 1.2 NA 2.0 Naphtha (C ₅ -350 ℉) 6.8 2.2 6.8 3.3 11.8 5.0 11.9  Distillate (350-700 ° F.) 28.0 23.5 29.0 24.4 32.0 33.0 39.4  Bottom water (700 ℉ +) 63.8 74.2 63.3 71.5 57.9 62.0 47.6 Total liquid product  -API specific gravity 31.5 28.0 31.6 28.6 34 31.3 34.3  Sulfur, ppm 486 800 302 400 108 60 50  Nitrogen, ppm 37 85 24 35 13 10 8  Flow point, ℃ - - 13 7 15 5 -8

As can be seen above, substitution of 50% of the hydrocracking catalyst with a sour service hydroisomerization catalyst seems to reduce the 650 ° F. + conversion under certain conditions. However, at a constant conversion rate, distillate yield is significantly increased and the total liquid product pour point is greatly reduced. Correspondingly, naphtha yield is reduced, and the pour point of the distillate and bottom product seems to decrease. The bottoms yield is similar at lower conversions but lower at higher conversions. Yields are shown in FIGS. 1, 2 and 3.

Example  3: sour  service Hydrolysis Process evaluation of hydrogen isomerization

The total liquid product in Example 2 was collected and distilled and analyzed for fuel and lubricant yield and properties. See Tables 4 to 6 and FIG. 4.

For Comparative Example-Hydrolysis (R2) after Hydrotreatment (R1) Lubricant Characteristics Comparative example Comparative example Comparative example Comparative example Comparative example 700 ℉ + conversion rate, weight% 35 40 43 63 73 Lube oil pour point, ℃ 41 37 45 37 39 700 ° F yield, wt% 65 60 57 37 27

In the case of the present invention-lubricant properties for HDC / waxing (R2) after hydrotreating (R1) Lubricant Characteristics HDT / HDC / HI HDT / HDC / HI HDT / HDC / HI HDT / HDC / HI HDT / HDC / HI HDT / HDC / HI HDT / HDC / HI 700 ℉ + conversion rate, weight% 26 38 39 50 51 63 79 Lube oil pour point, ℃ 21 12 7 -3 4 -7 -12 Lubricant viscosity at 100 ° C., cSt 7.2 4.9 Lubricant V.I. 105 112 700 ° F yield, wt% 74 62 61 50 49 37 21 Lubricant% Saturates ^* 70 84 ^* % Saturates (% by weight) = [1- (700 ° F + total aromatics in moles (g / g) * calculated molecular weight)] * 100
Wherein the molecular weight is calculated based on the kinematic viscosity at 100 ° C. and 40 ° C. of the 700 ° F. + lubricant

Diesel fuel properties for HDC / waxing (R2) after hydrotreatment (R1) in the present invention with comparative case (HT / HDC) Diesel characteristics HDT / HDC / HI Comparative example Diesel cloud point, ℃ -30.3 3.1 Calculated Setane Index of Diesel ^* 51.7 50.4 Diesel API 34 32 Diesel yield, weight percent 45 32 700 ℉ + conversion rate, weight% 50 40 ^* Cetane index is calculated according to ASTMD976

Hydrolysis (HDC) and hydroisomerization (HI) after integrated hydrotreating (HDT) showed improved diesel yield and diesel low temperature properties compared to the comparative example. In addition, the diesel characteristics for the integrated process were the same as in the comparative example, as indicated by the calculated stein index.

Example  3: semi-high pressure Sweet ( Semi - Sweet ) service Hydrolytic  Process evaluation

In order to assess the benefits of the intermediate removal of NH ₃ and H ₂ S after the hydrotreating zone, the hydrotreated MVGO from R 1 was stripped to remove NH ₃ , followed by the procedure of R 2. At constant T and LHSV, a significant increase in conversion and yield was observed.

R1-R2 Direct Cascade 650+ conversion rate weight% 50 54 60 65 LPG weight% 3 3 4 4 naphtha weight% 11 13 16 20 Distillate weight% 44 46 46 47 Bottoms weight% 42 38 34 29 R1-NH ₃ stripping-R2 650+ conversion rate weight% 53 63 78 87 LPG weight% 3 3 4 4 naphtha weight% 11 15 21 26 Distillate weight% 49 51 52 52 Bottoms weight% 37 31 23 18

All patents and patent applications, test procedures (e.g., ASTM methods, UL methods, etc.) and other documents cited herein are referenced in detail to the extent that the above description is consistent with the present invention and all areas in which it is permitted to be incorporated herein. have.

Where numerical lower limits and numerical upper limits are listed herein, ranges from any lower limit to any upper limit are considered. While exemplary embodiments of the invention have been described in particular, it will be understood that various other modifications can be made apparent and readily by those skilled in the art without departing from the spirit and scope of the invention. Accordingly, there is no intention to limit the claims appended hereto as to the embodiments and inventions described herein, and includes all features recognized as equivalent by one of ordinary skill in the art to which this invention belongs. Therefore, it is to be construed that the claims include all patentable novel features caused by the invention.

The invention has been described above in connection with a number of aspects and specific embodiments. Many modifications may be made to those skilled in the art in light of the above detailed description. All such obvious modifications also fall within the intended full scope of the appended claims.

Claims (95)

Contacting the hydrotreated feedstock and hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent,
Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and
Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions
Including but not limited to:
The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,
The hydrolysis catalyst comprises a zeolite Y-based catalyst,
The dewaxing catalyst comprises at least one non- dealuminated one-dimensional 10-member ring pore zeolite, at least one Group VIII metal and at least one low surface area metal oxide refractory binder,
How to make naphtha fuel, diesel fuel and lube basestock The method of claim 1,
Hydrotreating the entire hydrotreated, hydrolyzed and dewaxed effluent under effective hydrotreating conditions. The method of claim 2,
Fractionally distilling the entire hydrotreated, hydrolyzed, dewaxed, hydrotreated effluent to produce at least a lubricating oil basestock fraction; And further dewaxing the lubricating oil basestock fraction. The method of claim 3, wherein
The further dewaxing of the lubricant basestock fraction comprises at least one of solvent dewaxing the lubricant basestock fraction and catalytic dewaxing the lubricant basestock fraction. The method of claim 3, wherein
A method in which the dewaxed lubricant basestock is hydrotreated and stripped under reduced pressure under effective hydrofinishing conditions. The method of claim 1,
Wherein the hydrogen gas is selected from hydrotreated gas effluent, clean hydrogen gas, recycle gas, and combinations thereof. The method of claim 1,
Wherein the hydrotreated feedstock is cascaded into the hydrocracking step without separation. The method of claim 1,
Wherein the dewaxing catalyst comprises a molecular sieve having a SiO ₂ : Al ₂ O ₃ ratio of 200: 1 to 30: 1 and comprising a skeletal Al ₂ O ₃ content of 0.1% to 3.33% by weight. The method of claim 8,
The molecular sieve is EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, ZSM-22, EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof How water is. The method of claim 8,
The molecular sieve is EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof. The method of claim 8,
The molecular sieve is ZSM-48, ZSM-23 or a combination thereof. The method of claim 8,
The molecular sieve is ZSM-48. The method of claim 1,
The metal oxide refractory binder has a surface area of 100 m ² / g or less. The method of claim 1,
The metal oxide refractory binder has a surface area of 80 m ² / g or less. The method of claim 1,
The metal oxide refractory binder has a surface area of 70 m ² / g or less. The method of claim 1,
Wherein the surface area of the micropores to the total surface area of the dewaxing catalyst is at least 25%, and the total surface area is the sum of the surface area of the outer zeolite and the surface area of the binder. The method of claim 1,
The metal oxide fire resistant binder is silica, alumina, titania, zirconia or silica-alumina. The method of claim 1,
And wherein the metal oxide refractory binder further comprises a second metal oxide refractory binder different from the first metal oxide refractory binder. The method of claim 18,
The second metal oxide is silica, alumina, titania, zirconia or silica-alumina. The method of claim 1,
The dewaxing catalyst comprises from 0.1 to 5% by weight of platinum. The method of claim 1,
Wherein the hydrocracking and dewaxing steps are carried out in a single reactor. The method of claim 1,
Wherein the hydrocracking and dewaxing steps are carried out in two or more reactors in series. The method of claim 2,
Hydrocracking, dewaxing and a second hydrotreating step are carried out in a single reactor. The method of claim 2,
Wherein the hydrocracking, dewaxing and second hydrotreating steps are carried out in two or more reactors in series. The method of claim 2,
Wherein the first hydrotreatment, hydrocracking, dewaxing and second hydrotreating step are carried out in a single reactor. The method of claim 2,
Wherein the first hydrotreating, hydrocracking, dewaxing, and second hydrotreating steps are carried out in two or more reactors in series. Contacting the hydrotreated feedstock and the hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent, wherein prior to said contacting, the effluent from the hydrotreating step is fed to one or more high-pressure separators to provide hydrogen Separating the gaseous fraction of the hydrotreated effluent from the liquid fraction of the treated effluent,
Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and
Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions
Including but not limited to:
The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,
The hydrolysis catalyst comprises a zeolite Y-based catalyst,
Wherein the dewaxing catalyst comprises at least one non- dealuminated one-dimensional 10 membered ring pore zeolite, at least one Group VIII metal and at least one low surface area metal oxide refractory binder,
How to make naphtha fuel, diesel fuel and lubricant basestock. The method of claim 27,
The hydrotreated effluent after separation comprises dissolved H ₂ S and any organic sulfur. The method of claim 27,
Wherein the hydrotreated effluent is separated with the hydrogen containing gas after separation. 30. The method of claim 29,
Wherein the hydrogen containing gas comprises H ₂ S. The method of claim 27,
Wherein the hydrogen gas is selected from hydrotreated gas effluent, clean hydrogen gas, recycle gas, and combinations thereof. The method of claim 27,
Hydrotreating the entire hydrotreated, hydrolyzed and dewaxed effluent under effective hydrotreating conditions. 33. The method of claim 32,
Fractionally distilling the entire hydrotreated, hydrolyzed, dewaxed, hydrotreated effluent to produce at least a lubricating oil basestock fraction; And further dewaxing the lubricating oil basestock fraction. 34. The method of claim 33,
The further dewaxing of the lubricant basestock fraction comprises one or more of solvent dewaxing the lubricant basestock fraction and / or contact dewaxing the lubricant basestock fraction. 34. The method of claim 33,
Additionally the dewaxed lubricant basestock is hydrotreated and decompressed stripped under effective hydrofinishing conditions. The method of claim 27,
Wherein the dewaxing catalyst comprises a molecular sieve having a SiO ₂ : Al ₂ O ₃ ratio of 200: 1 to 30: 1 and comprising a skeletal Al ₂ O ₃ content of 0.1% to 3.33% by weight. The method of claim 36,
The molecular sieve is EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, ZSM-22, EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof How is water. The method of claim 36,
The molecular sieve is EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof. The method of claim 36,
The molecular sieve is ZSM-48, ZSM-23 or a combination thereof. The method of claim 36,
The molecular sieve is ZSM-48. The method of claim 27,
The metal oxide refractory binder has a surface area of 100 m ² / g or less. The method of claim 27,
The metal oxide refractory binder has a surface area of 80 m ² / g or less. The method of claim 27,
The metal oxide refractory binder has a surface area of 70 m ² / g or less. The method of claim 27,
Wherein the surface area of the micropores to the total surface area of the dewaxing catalyst is at least 25%, and the total surface area is the sum of the surface area of the outer zeolite and the surface area of the binder. The method of claim 27,
The metal oxide fire resistant binder is silica, alumina, titania, zirconia or silica-alumina. The method of claim 27,
And wherein the metal oxide refractory binder further comprises a second metal oxide refractory binder different from the first metal oxide refractory binder. 47. The method of claim 46,
And the second metal oxide fire resistant binder is silica, alumina, titania, zirconia or silica-alumina. The method of claim 27,
The dewaxing catalyst comprises from 0.1 to 5% by weight of platinum. The method of claim 27,
Wherein the hydrocracking and dewaxing steps are carried out in a single reactor. The method of claim 27,
Wherein the hydrocracking and dewaxing steps are carried out in two or more reactors in series. 33. The method of claim 32,
Hydrocracking, dewaxing and a second hydrotreating step are carried out in a single reactor. 33. The method of claim 32,
Wherein the hydrocracking, dewaxing and second hydrotreating steps are carried out in two or more reactors in series. 33. The method of claim 32,
Wherein the first hydrotreating, hydrocracking, dewaxing, and second hydrotreating steps are carried out in two or more reactors in series. Contacting the hydrotreated feedstock and hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent,
Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and
Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions
Including but not limited to:
The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,
The hydrolysis catalyst comprises a zeolite Y-based catalyst,
The dewaxing catalyst comprises at least one non- dealuminated one-dimensional 10 membered ring pore zeolite and at least one Group VIII metal,
How to make naphtha fuel, diesel fuel and lubricant basestock. 55. The method of claim 54,
Hydrotreating the entire hydrotreated, hydrolyzed and dewaxed effluent under effective hydrotreating conditions. 56. The method of claim 55,
Fractionally distilling the hydrotreated, hydrolyzed, dewaxed, and hydrotreated effluent to produce at least a lubricating oil basestock fraction; And further dewaxing the lubricating oil basestock fraction. 57. The method of claim 56,
The further dewaxing of the lubricant basestock fraction comprises one or more of solvent dewaxing the lubricant basestock fraction and / or contact dewaxing the lubricant basestock fraction. 57. The method of claim 56,
A method in which the dewaxed lubricant basestock is hydrofinishing and pressure stripping under effective hydrofinishing conditions. 55. The method of claim 54,
Wherein the hydrogen gas is selected from hydrotreated gas effluent, clean hydrogen gas, recycle gas, and combinations thereof. 55. The method of claim 54,
Wherein the hydrotreated feedstock is cascaded into the hydrocracking step without separation. 55. The method of claim 54,
Wherein the dewaxing catalyst comprises a molecular sieve having a SiO ₂ : Al ₂ O ₃ ratio of 200: 1 to 30: 1 and comprising a skeletal Al ₂ O ₃ content of 0.1% to 3.33% by weight. 62. The method of claim 61,
The molecular sieve is EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, ZSM-22, EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof How is water. 62. The method of claim 61,
The molecular sieve is EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof. 62. The method of claim 61,
The molecular sieve is ZSM-48, ZSM-23 or a combination thereof. 62. The method of claim 61,
The molecular sieve is ZSM-48. 55. The method of claim 54,
The dewaxing catalyst is self-bonding and does not contain a binder. 55. The method of claim 54,
The surface area of the micropores to the total surface area of the dewaxing catalyst is at least 25%, and the total surface area is the same as that of the outer zeolite. 55. The method of claim 54,
The dewaxing catalyst comprises from 0.1 to 5% by weight of platinum. 55. The method of claim 54,
Wherein the hydrocracking and dewaxing steps are carried out in a single reactor. 55. The method of claim 54,
Wherein the hydrocracking and dewaxing steps are carried out in two or more reactors in series. 56. The method of claim 55,
Hydrocracking, dewaxing and a second hydrotreating step are carried out in a single reactor. 56. The method of claim 55,
Wherein the hydrocracking, dewaxing and second hydrotreating steps are carried out in two or more reactors in series. 56. The method of claim 55,
Wherein the first hydrotreatment, hydrocracking, dewaxing and second hydrotreating step are carried out in a single reactor. 56. The method of claim 55,
Wherein the first hydrotreating, hydrocracking, dewaxing, and second hydrotreating steps are carried out in two or more reactors in series. Contacting the hydrotreated feedstock and the hydrogen containing gas with a hydrocracking catalyst under effective hydrocracking conditions to produce a hydrocracked effluent, wherein prior to said contacting, the effluent from the hydrotreating step is fed to one or more high-pressure separators to provide hydrogen Separating the gaseous fraction of the hydrotreated effluent from the liquid fraction of the treated effluent,
Cascading the entire hydrocracked effluent to a contact dewaxing stage without separation, and
Dewaxing the entire hydrocracked effluent under effective contact dewaxing conditions
Including but not limited to:
The total sulfur in liquid and gaseous form supplied to the dewaxing stage is greater than 1000 ppm by weight sulfur, based on the hydrotreated feedstock,
The hydrolysis catalyst comprises a zeolite Y-based catalyst,
Wherein the dewaxing catalyst comprises at least one non- dealuminated one-dimensional 10 membered ring pore zeolite and at least one Group VIII metal,
How to make naphtha fuel, diesel fuel and lubricant basestock. 78. The method of claim 75,
The hydrotreated effluent after separation comprises dissolved H ₂ S and any organic sulfur. 78. The method of claim 75,
Wherein the hydrotreated effluent is separated with the hydrogen containing gas after separation. 78. The method of claim 77,
Wherein the hydrogen containing gas comprises H ₂ S. 78. The method of claim 75,
Wherein the hydrogen gas is selected from hydrotreated gas effluent, clean hydrogen gas, recycle gas, and combinations thereof. 78. The method of claim 75,
Hydrotreating the entire hydrotreated, hydrolyzed and dewaxed effluent under effective hydrotreating conditions. 79. The method of claim 80,
Fractionally distilling the entire hydrotreated, hydrolyzed, dewaxed, hydrotreated effluent to produce at least a lubricating oil basestock fraction; And further dewaxing the lubricating oil basestock fraction. 83. The method of claim 81,
The further dewaxing of the lubricant basestock fraction comprises one or more of solvent dewaxing the lubricant basestock fraction and / or contact dewaxing the lubricant basestock fraction. 83. The method of claim 81,
Additionally the dewaxed lubricant basestock is hydrotreated and decompressed stripped under effective hydrofinishing conditions. 78. The method of claim 75,
Wherein the dewaxing catalyst comprises a molecular sieve having a SiO ₂ : Al ₂ O ₃ ratio of 200: 1 to 30: 1 and comprising a skeletal Al ₂ O ₃ content of 0.1% to 3.33% by weight. 85. The method of claim 84,
The molecular sieve is EU-1, ZSM-35, ZSM-11, ZSM-57, NU-87, ZSM-22, EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof How is water. 85. The method of claim 84,
The molecular sieve is EU-2, EU-11, ZBM-30, ZSM-48, ZSM-23 or a combination thereof. 85. The method of claim 84,
The molecular sieve is ZSM-48, ZSM-23 or a combination thereof. 85. The method of claim 84,
The molecular sieve is ZSM-48. 78. The method of claim 75,
The surface area of the micropores to the total surface area of the dewaxing catalyst is at least 25%, and the total surface area is the same as that of the outer zeolite. 78. The method of claim 75,
The dewaxing catalyst comprises from 0.1 to 5% by weight of platinum. 78. The method of claim 75,
Wherein the hydrocracking and dewaxing steps are carried out in a single reactor. 78. The method of claim 75,
Wherein the hydrocracking and dewaxing steps are carried out in two or more reactors in series. 79. The method of claim 80,
Hydrocracking, dewaxing and a second hydrotreating step are carried out in a single reactor. 79. The method of claim 80,
Wherein the hydrocracking, dewaxing and second hydrotreating steps are carried out in two or more reactors in series. 79. The method of claim 80,
Wherein the first hydrotreating, hydrocracking, dewaxing, and second hydrotreating steps are carried out in two or more reactors in series.

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