JP3119489B2 - Integrated hydroprocessing with separation and recycling - Google Patents

Description

DETAILED DESCRIPTION OF THE INVENTION The present invention relates to the refining of petroleum products, in particular to the integrated hydroproceing of middle distillates and lubricating oil range boiling materials.
ssing) process (including hydrocracking and subsequent dewaxing using shape-selective or hydroisomerization catalysts), the resid streams of each stage (and optionally other streams) are separated from one another during processing. The method relates to a method of holding the information. The residual oil fraction from the dewaxing step can be recycled back to the hydrocracking step for further processing or used as a lubricant feedstock.

Conventional petroleum hydrocracking and catalytic dewaxing have been described by fundamentally different process chemistries. Naphthenes and aromatics are more easily hydrocracked than paraffins due to their molecular structure. The energy required to saturate and open the aromatic compound is less than the energy required to break the paraffinic bonds. As a result, in conventional high conversion recycle hydrocracking, the recycled oil stream becomes paraffin-rich and is characterized by a very high pour point, sometimes even exceeding 38 ° C (100 ° F) . "Shape-selective" catalysts limit the decomposition of the ring structure,
Normal paraffins are converted preferentially, achieving substantial pour points, and may result in catalytic dewaxing, especially molecular weight reduction in the manufacture of lubricants. Certain types of dewaxing catalysts are designed to convert linear paraffins by isomerization in combination with cracking, so that they are selective for middle distillate products, such as kerosene and diesel oil. is there.

The present invention discloses an integrated hydroprocessing system using both hydrocracking and catalytic dewaxing. Unconverted hydrocracker bottoms material is dewaxed in a dewaxing stage. The unconverted material from the dewaxing step is subsequently recycled back to the hydrocracking step. Feed to catalytic dewaxing stage-hydrocracked residual oil fraction-
Is generally rich in linear paraffin. These linear paraffins can be isomerized in the dewaxing step to produce paraffins with more branched chains suitable as gasoline and middle distillate components.

In the dewaxing step, a single catalyst or a two-catalyst system can be used. Large pore catalysts, such as zeolite B, can be used to some extent for isomerizing or cracking linear paraffins, depending on the catalytic activity. Shape selective intermediate pore size catalysts, such as ZSM-5, can be used to degrade linear paraffins.

Integrated hydrotreating processes using both hydrocracking and catalytic dewaxing steps are well known in the purification arts. U.S. Pat.Nos. 4,851,109 and 4,764,266,
Disclosed is an integrated hydroprocessing method for hydrocracking high boiling feeds, such as gas oil and catalytic cracked cycle oil, with an aromatic-selective hydrocracking catalyst to produce naphtha and middle distillate range products doing. Unconverted material can be recycled to the hydrocracking process,
Alternatively, it can be sent to a dewaxing step and treated with a large pore catalyst having hydrogenation-dehydrogenation functionality, especially zeolite beta, to produce further kerosene and distillate range products. Although this patent employs the recycling of the hydrocracker bottoms product, the unconverted material from the dewaxing step is not recycled to the hydrocracking process. further,
These methods describe the use of an atmospheric distillation unit having baffles in a flash zone to separate resid materials from two different process streams, as described in one aspect of the invention. Not.

PCT application WO 95/10578 discloses a process for producing middle distillate products by dewaxing hydrocarbonaceous materials having a boiling point of 343 ° C. (650 ° F.) or higher. The feed is hydrocracked by a catalyst comprising a large pore size zeolite, such as zeolite Y, and a metal hydride component selected from Groups VIB or VIII of the Periodic Table. The entire effluent of the hydrocracking zone is then dewaxed in the presence of hydrogen over a catalyst comprising a medium pore size zeolite such as ZSM-5. In the present invention, only the residual oil substance of the hydrocracker distillate is dewaxed.

EP 189648 A1 discloses a method for hydrotreating a heavy gas oil feed to maximize kerosene and distillate production. FIG. 5 of this patent shows the hydrocracking of the feed followed by catalytic dewaxing. In the present invention, the entire distillate of the hydrocracking unit is dewaxed, and not the residual oil substance alone as in the present invention. Furthermore, nothing is taught about recycling to the first hydrocracking stage. Two different hydrocracking steps are used.

U.S. Patent No. 5,053,117 (Kyan et. Al.) Includes MPHC
(Moderate Pressure Hydrocracker Bottoms) Method
Distillate Dewaxing Process) A method for producing naphtha and light oil under conditions, or less severe MLDW
(Mobil Lude Dewaxing process) A method for producing a lubricating oil under conditions is disclosed. The feed to the dewaxing unit can improve the octane number of the naphtha product by mixing gas oil with a high nitrogen content. Recycling of the dewaxer resid material to the hydrocracker is not taught. The octane number of gasoline can be improved by maximizing olefin production. This can be achieved by using MDDW catalysts and conditions with minimal hydrogenation.

Various patents describe a general method for hydrotreating feeds, such as heavy gas oils (or other hydrocarbon feeds boiling above 343 ° C. (650 ° F.), and then dewaxing to produce lubricating oils. Is explained. U.S. Patent No. 4,414,097
Discloses a method of hydrocracking a hydrocarbon feedstock having a boiling point of 343 ° C. (650 ° F.) or higher and then dewaxing with a catalyst containing ZSM-23.

U.S. Patent Nos. 4,282,271 (Garwood et al.) And 4,28,271.
No. 3,272 (Garwood et al.) Discloses hydrotreating a lubricating oil by hydrocracking a feedstock boiling above 343 ° C. (650 ° F.). The hydrocrackate is then dewaxed. The effluent from the dewaxing unit is hydrotreated but not recycled to the hydrocracking step.

The present invention relates to the refining of petroleum products, and in particular to an integrated hydrotreating process (including a hydrocracking step followed by a step of dewaxing unconverted hydrocracker resid material). The resid stream (and possibly other streams in each step) are kept independent of each other. Dewaxing is carried out using a hydroisomerization catalyst or a shape-selective catalyst or both. In one embodiment, using a baffle in the flash zone of the fractionator,
The resid streams are separated from one another. Alternatively, the effluent from the hydrocracking step can be treated independently from the effluent from the dewaxing step. The residual oil fraction from the dewaxing step can be recycled back to the hydrocracking step for further processing or used as a lubricating oil feedstock.

In the integrated hydrotreating method of the present invention, the recycling process can achieve the maximum yield at the minimum cost. The process is versatile and can be performed at pressures as high as the hydrocracking (high pressure dewaxing) step, or at substantially lower pressures resulting in less hydrogen consuming conversions. it can. Further, the product can be removed from a single fractionator or from two or more fractionators. In this way, the product obtained after dewaxing can be separated from the product obtained after the hydrocracking step. Although a single fractionator can be used to recover common products, at least a portion of the dewaxed distillate can be recycled to the hydrocracker, so unconverted from the two reaction stages It is desirable to keep the resid product separately. In one embodiment, the reactor products from the hydrocracking and dewaxing stages can be stripped separately. The stripped stream mixture can then be sent to a debutanizer. LPG can be recovered at minimum cost. The liquid product stream of the reactor is charged separately to the fractionator (see FIGS. 1 and 3).

In another embodiment, a make-up compressor can be used to provide a one-pass hydrogen stream during the compression stage. The dewaxing stage can be operated at a low gas to oil ratio and low pressure. The outlet hydrogen of the dewaxing unit is then compressed and fed to the hydrocracking step for make-up (replenishment), eliminating the need for a separate recycle hydrogen compressor in the dewaxing step (see FIG. 3).

By dewaxing substances having a boiling point in the middle distillate range by isomerization, the cold flow properties of the product, such as the pour point to diesel, the cloud point and the freezing point of jet fuel, can be improved. By improving the paraffinic properties of kerosene and diesel products in the present invention, the combustion characteristics of middle distillates are also improved.

FIG. 1 illustrates the integrated hydrocracking and dewaxing steps of the present invention. Although not shown, the hydrocracking stage may include two hydrotreating-hydrocracker catalytic systems. In the subsequent hydrocracking, the feed is separated into the respective products in a fractionator by atmospheric distillation. The resid material of the hydrocracking step is subsequently dewaxed.
The fractionator of FIG. 1 shows baffles in the flash zone to keep the resid fractions from the hydrocracking and dewaxing stages separate. In FIG. 2, the hydrocracked effluent and the dewaxed effluent are treated separately by using separate fractionators and debutane units. This is desirable when dewaxing or isomerization is performed at high pressure. Figures 1 and 3 show that the reactor products from the two stages are separately stripped. The stripped stream mixture is then debutanized. FIG. 3 shows the hydrocracking and dewaxing stages operating at approximately equivalent pressure levels.

Feeds Feedstocks used in the present invention are generally characterized as high boiling feeds of mineral oil origin, but feeds of other origins, such as synthetic oil production processes such as Fi
Feeds from the scher-Tropsch synthesis method or other synthesis methods, such as methanol conversion methods, can also be used. Fractions from unconverted sources, such as shale oil and tar sands, can also be processed by the integrated hydroprocessing techniques of the present invention. Generally, the feed has a relatively high boiling point, usually above 205 ° C. (400 ° F.), such as above 230 ° C. (450 ° F.), and often above 315 ° C. (600 ° F.). And often has an initial boiling point of 345 ° C (650 ° F) or higher. The boiling point characteristics, especially the end point, of the feed are determined by the products required. If a significant amount of lubricating oil is to be processed, the feed must itself contain significant amounts of components above the lubricating oil boiling range, typically 650 ° F (345 ° C). Thus, if lubricating oil production is desired, the feed will generally be a gas oil, for example, with an endpoint typically at 565 ° C (1050 ° F).
A high boiling distillate feed that is: However, it is not intended to rule out the presence of significant amounts of non-distillable residues in the feed that can be processed. Deasphalted oil (oil obtained from a propane deasphalting unit and suitable for the manufacture of bright stock) can also be used as feed. Typical feeds that can be processed include gas oils, such as heavy coker gas oils, vacuum gas oils, vacuum crude oils and atmospheric gas oils. In contrast, when the process of the present invention is used primarily for the production of middle distillates and naphtha, especially jet fuel, the feed may have a lower end point because it does not need to tolerate higher boiling components. Therefore, to produce jet fuel, light cycle oil (LCO), heavy cycle oil (HCO),
Catalytic cracking cycle oils, including clarified slurry oil (CSO) and other catalytic cracking products, are particularly useful feed sources for the process of the present invention. Cycle oils from catalytic cracking processes are generally between 205-400 ° C (400 ° F-750 ° F)
, But light cycle oils may have a lower endpoint, for example, an endpoint of 600 ° F or 650 ° F (315 ° C or 345 ° C). Heavy cycle oils are higher, for example
It may have an initial boiling point (IBP) of 260 ° C (500 ° F). The cycle oil has a relatively high aromatics content which makes it very suitable for processing in the first hydrocracking step of the integrated process sequence of the present invention, and furthermore, such a refractory material The lower required level makes it a very preferred substance for processing in the method of the invention.

The feeds used in the process of the invention generally contain a relatively high proportion of aromatics, especially polycyclic aromatics, due to their high boiling points, but also high amounts of high boiling paraffins and cycloparaffins. Exists in. When the feed is subjected to catalytic cracking, the aromatics are substantially dealkylated and are less likely to be further cracked except in the case of a hydrogenation process as in the present invention. The process of the present invention comprises the steps of converting a refractory aromatic compound feed, such as one obtained from a catalytic cracking operation, into a high quality, high paraffinic distillate product having a low pour point, as well as a low pour point lubricating oil. It should be noted that it can be converted to

Generally, the aromatics content of the feed used in the process of the invention is at least 30% by weight, usually at least 40% by weight.
% By weight, and often at least 50% by weight. The balance is distributed between paraffin and naphthene, depending on the origin of the feed and its pretreatment. Catalytically cracked materials tend to have higher aromatics content than other feeds, and in some cases aromatics content.
It may exceed 60% by weight. Prior to hydrocracking, the feed may be hydrotreated to separate contaminants.

Tables 1 to 4 show the composition of a typical vacuum gas oil (VGO).
Tables 2 and 3 show two hydrotreated (HDP) gas oils, and Tables 5 and 6 show two typical FCC LCOs.
Shows a feed.

Catalysts and Conditions-First Hydrocracking Stage The purpose of the hydrocracking performed in the first stage of the integrated hydrotreating process of the present invention is to carry out the second stage using a catalyst or shape selective dewaxing catalyst. To form a feed having a relatively high paraffin concentration. Therefore, the catalyst used in the first stage hydrocracking is a catalyst having selectivity for aromatic compounds over paraffin, but the hydrocracking of paraffin is not hindered. Thus, at this stage, conventional hydrocracking catalysts using large pore size solid materials with acidic functionality combined with a hydrogenation function can be used. The use of large pore size materials allows the bulky polycyclic aromatic components in the feed to have access to the internal pore structure of the catalyst where the characteristic cracking reactions primarily occur, thus providing a high boiling feed, e.g. It is considered important for the hydroprocessing of gas oil. Degradation occurs to a limited extent on the outer surfaces of the catalyst particles, but a major portion of the catalyst surface area resides inside the particles. It is important that this internal pore structure be accessible to the components of the feed that undergo degradation. Therefore,
The acidic functionality of the first stage hydrocracking catalyst may be determined by a large pore size amorphous material, such as alumina, silica-alumina or silica, or by a large pore size crystalline material, preferably a large pore size alumina silicate zeolite, For example, zeolite X, zeolite Y, ZSM-3, ZSM-18, ZSM
-20, MCM-22 or zeolite beta. The zeolites are more stable, resistant to decomposition and thus loss of acid functionality, even under the influence of various cationic types or other types, preferably the hydrothermal conditions encountered during the hydrocracking reaction. It can be used in a high form. Therefore, a form with improved stability,
For example rare earth exchanged large pore zeolites such as REX and REY
And so-called ultrastable zeolite Y (US
Y) and high silica zeolites such as dealuminated Y or dealuminated mordenite are preferred.
Particularly preferred in the present invention are dealuminated USY catalysts having a backbone silica / alumina ratio of 50: 1 or more.

Zeolite ZSM-3 is disclosed in U.S. Pat.No. 3,415,736, zeolite ZSM-18 is disclosed in U.S. Pat.No. 3,950,496, and zeolite ZSM-20 is disclosed in U.S. Pat.
No. 72,983, whose description, properties and preparation can be referred to for these zeolites.

The crystalline component can be incorporated into an amorphous matrix for higher physical strength, which matrix can itself be catalytically active. For example, amorphous alumina or silica having acidic functionality
It can be alumina or clay, which participates in the decomposition reaction catalyzed by the resulting acid. The matrix will generally make up 40-80%, more usually 50-75%, by weight of the catalyst.

The hydrogenation is effected in a conventional manner by means of the transition metal components of the base or noble metals, generally groups VIA and VIIIA of the periodic table.
It can be given by using a tribe. Such metals include nickel, cobalt, tungsten, molybdenum, palladium or platinum, with the base metals nickel, tungsten, cobalt and molybdenum being preferred. Such combinations of nickel-tungsten, nickel-cobalt and cobalt-molybdenum are particularly useful, especially when the feed has selectively high levels of contaminants.

Under low to medium severity conditions, it is selective for aromatics and operates at 165 ° C. (330 ° C.), containing a relatively high content of naphthenes, by operating with a catalyst intended for naphtha. F) A naphtha product having the following boiling point is formed as a major part and is therefore of value for reforming to produce high octane gasoline. In addition, small to moderate middle distillates are formed with high boiling fractions.

If the hydrocracking step is operated in distillate selection mode at low to moderate hydrogen pressures and low to moderate severity conditions, as described, for example, in U.S. Pat. The separated fraction is a major middle distillate, typically 165 ° C to 345 ° C (330 ° F to 650 ° F)
And a distillate containing a lower proportion of naphtha. When a distillate-selective catalyst is used, middle distillates can be a major part of the product, as disclosed in US Pat. No. 4,820,402. Middle distillate fractions can have very aromatic character when operated at relatively low pressures. The aromatics content in the product is generally lower when higher pressures are used. The middle distillate fraction is generally unsuitable for direct use as jet fuel or diesel when the aromatics are higher, but when mixed with other components, for example those with a higher paraffinicity,
It can be used as a compounding component in diesel oil.
It can be used as such a fuel oil, or can be used in combination with other components having a higher boiling point.
This middle distillate fraction, however, has a relatively low sulfur content and generally satisfies product specifications for use as light fuels, such as domestic fuel oils. The unconverted or resid fraction from the low to medium pressure first stage hydrocracking is generally waxy and has very low sulfur and nitrogen contents. Although this product is valuable and may be sold as a low sulfur heavy fuel oil, processing it in the isomerization or dewaxing step significantly improves distillate yield and quality. . The paraffin remaining in the hydrocracker bottoms has a high isoparaffin content in the second stage and is selectively converted to lower boiling products.

When operating the hydrocracking process in naphtha selection mode, moderate to high severity conditions are used. Thus, a relatively acidic hydrocracking catalyst or a combination of a hydrotreating / hydrocracking catalyst and a propensity for naphtha may be used at relatively high temperatures, e.g., 400-450 ° C (750-850 ° F).
Tend to be used at temperatures up to the upper limit of the hydrocracking range. Temperatures of this order can be easily achieved by controlling the quenching (cooling) of hydrogen in the reactor, but the severity can also be adjusted by appropriately controlling the space velocity. To promote the saturation of the aromatics in the feed, a high hydrogen partial pressure, typically 70
00kpa (1000psig) or more, often 10,000kpa (1435psi)
g) Use of the above pressures is common and preferred for naphtha mode operation. Hourly space velocity (LHSV) can be typically a 2 hr ^-1, a 1hr ^-1 following values in the case of desired lower operating temperature. The hydrogen circulation flow rate is selected to maintain catalytic activity at the selected temperature, and is generally 35
0n.ll ^-1 (1966SCF / Bbl), often 500-1
000n.ll ^-1 (2810-5620SCF / Bbl). A hydrocracking process selective to naphtha is very effective in reducing sulfur and nitrogen levels in the unconverted resid fraction, so that the process has low levels of sulfur and nitrogen and Characterized by high levels of waxy paraffin. These paraffins are then selectively isomerized and hydrocracked by a second stage operation to produce high quality distillate and kerosene products.

The change in product distribution that occurs when using selective hydrocracking for naphtha in the first stage can be expressed as: The feed containing paraffins, naphthenes and aromatics has significant naphthene / It is converted to a full range product with aromatic naphtha. Aromatics in the feed are reduced by the aromatics selectivity of the first stage catalyst. When the paraffinic resid fraction is subjected to a second stage hydrotreatment using a zeolite beta-based catalyst, paraffins are preferentially converted to low boiling isoparaffins in the middle distillate boiling range, producing small amounts of naphtha. The unconverted fraction is a low pour point product because the paraffin therein is isomerized to isoparaffin.

The operation of the naphtha mode preferably employs a more acidic catalyst, higher severity and higher hydrogen pressure, but as in many purification processes, the exact conditions used will depend on the nature of the feed. Selected. A single hydrocracking catalyst can be used, or a hydrotreating-hydrocracking catalyst system can be used. When achieving the desired medium to high severity conditions suitable for naphtha production, the temperature and space velocity (LHS
V) are interconnected. Thus, temperatures as low as 250 ° C (480 ° F) can be used, but 315 ° C (600 ° F).
Higher temperatures above F) are more common. When using a hydroprocessing / hydrocracking system, the operating temperature is 650-700 ゜
F range. The upper temperature limit is typically 450 ° C (850 ° F), with temperatures of 400-450 ° C (750-850 ° F) providing suitable harsh levels at higher space velocities, thereby accelerating catalyst aging Higher throughput is achieved at the expense of being done. Space velocities (LHSV) are typically less than 2 hr ^{-1 and} less than ¹ hr ^-1 at lower temperatures below 315 ° C. (600 ° F.). Operation in the naphtha mode preferably employs a higher hydrogen partial pressure, which promotes ring opening and decomposition to naphtha following aromatic saturation. Thus, a hydrogen partial pressure of at least 7000 kpa (1000 psig), preferably at least 10,000 kpa (1435 psig), is preferred. Fifteen
Higher pressures of 000 kpa (2160 psig) are highly preferred for aromatic saturation, but at the expense of increased hydrogen consumption.
The hydrogen circulation flow rate is generally at least 300 n.ll ^-1 (at least 1685 SCF / Bbl), preferably at least 500 n.
ll ^-1 (2800SCF / Bbl), usually 500-1000n.ll
^-1 (2800-5600SCF / Bbl). Conversion to naphtha and lower boiling range products (165 ° C.-) per pass depends on the nature and conditions of the feed, but is generally less than 5%.
-25% by weight, and the overall conversion per pass (i.e., conversion to lower boiling range products) is generally 15%.
It is in the range of 8080% by weight, often 20-60% by weight.

It is desirable to maximize the naphthenic / aromatic nature of the naphtha so that it is more amenable to subsequent modification. The resid product from the naphtha selective hydrocracking process is:
Since the catalyst used in this step attacks aromatics more than paraffin, it is concentrated in paraffin (relative to the feed). By treating the paraffin-rich resid fraction in the subsequent dewaxing step using a hydroisomerization catalyst, the overall yield of isoparaffinic liquid products, such as jet fuel and low pour point distillate, is higher. The resid fraction of the hydroisomerization step is relatively rich in naphthenes and aromatics and is recycled back to the hydrocracker. This operation results in higher naphtha treatment as well as higher naphthene / aromatic properties than operations that do not recycle or recycle the hydrocracker resid material.

The operation of hydrocracking in a mode aimed at distillate generally involves lower severity using lower acidity catalysts. Also in this case, 250 ° C to 450 ° C (480 ° F to 850 ° F)
, But higher space velocities up to 2 hr ^-1 may be suitable to achieve the desired severity. If an aromatic distillate product is acceptable, a low hydrogen partial pressure can be used as in the hydrocracking process described in U.S. Pat. Reference can be made to this document for a description of the preferred operating conditions. As described therein, a hydrogen pressure of 5250-7000 kpa (745-1000 psig) is sufficient, and a hydrogen circulation flow rate of 250-1000 n.ll ^-1 (1400
~ 5600SCF / Bbl), usually 300 ~ 800n.llC ^-1 (1685 ~ 45
00SCF / Bbl). Total conversion to 345 ° C .- (600 ° F.) product is generally less than 50%, typically 30-40%.

The change in composition resulting from the operation of the hydrocracking step in the low pressure distillate selection mode can be described as follows: the feed containing the high boiling fractions of paraffins, naphthenes and aromatics has a significant aromatic It is converted to a full range hydrocracking effluent containing a middle distillate with a family characteristic and relatively small amounts of naphtha. The unconverted fraction is again of paraffinic nature, due to the aromatic-selective nature of the catalyst used in this step. When contacted with the hydroisomerization catalyst in the second stage, paraffins are preferentially attacked to give lower boiling isoparaffins, mainly due to molecular weight reduction (decomposition), but less than just below the cut point. Some changes may also occur due to isomerization to lower boiling range isomers. Again, the middle distillate has a lower pour point due to its isoparaffinic content.

As described above, prior to the hydrocracking step, the contaminants, mainly sulfur and nitrogen, and any metal components present are removed prior to the hydrocracking step, thus providing a dual function hydrotreating / hydrocracking system. Can be formed. For this purpose, the hydrotreating conditions and the catalyst are chosen conventionally. Separation of the inorganic sulfur and nitrogen content between stages can be performed or can be omitted, as in the low pressure process described in US Pat. No. 4,435,275.

In the present invention, the hydrocracker residue is not recycled to the hydrocracker. All resid material is sent to a dewaxing or isomerization step. The resid material from this step is then recycled to the hydrocracker to obtain naphtha rich in aromatics and naphthenes.

If distillate-selective hydrocracking is used in the hydrocracking stage, the converted fractions are in the tail portion, e.g. relatively high boiling fractions, e.g.
The fraction between (440 ° F and 650 ° F) contains a relatively high proportion of paraffin. The cut point separating the portion fed to the second stage is thus adjusted, and if a hydroisomerization catalyst is used in the dewaxing step, these paraffins are hydrolyzed to a lower boiling range. Will shift. Thus, it may be desirable to provide a portion of the converted fraction to the second step along with the essentially unconverted fraction. The cut point may be as low as 200 ° C (390 ° F), but is generally at least 225 ° C (440 ° F) and often at least 315 ° C (600 ° F) , Generally 345 ° C (650 ° F).

Dewaxing Step Dewaxing is carried out by a hydroisomerization method, a shape-selective dewaxing method,
Alternatively, both can be used continuously. Both hydroisomerization and shape-selective dewaxing are described below.

A. Catalysts and Conditions for Performing Hydroisomerization 1. Catalysts This step is essentially a hydrogenation process using a combination of acid and hydrogenation functionalities based on large pore zeolites, for example, zeolite beta. Cracking or hydroisomerization step. Hydrogenation functionality can be provided by either base or noble metals. Nickel, tungsten, cobalt, molybdenum, palladium, platinum or combinations of such metals are preferred. Fluorinated nickel-tungsten combinations are particularly desirable. The acid functionality is a known zeolite and is known from U.S. Reissue Patent RE.
It is preferably provided by the zeolite beta described in US Pat. No. 28,341, to which reference can be made for a description of this zeolite, its preparation and its properties. A zeolite beta-based hydrotreating catalyst is disclosed in U.S. Pat.
Nos. 4,419,220, 4,501,926 and 4,518,485 as well as U.S. Patent Application Ser. These may be referred to for a description of possible zeolite betas. As described in these patents, the preferred form of zeolite beta for use in hydroprocessing is of the high silica type with a (structural) silica-alumina ratio of at least 30: 1. At least 5 to maximize paraffin isomerization at the expense of degradation
Silica-alumina ratios of 0: 1, preferably at least 100: 1, or higher or higher, such as 250: 1, 500: 1, can be used. Thus, by using a suitable silica: alumina ratio in the catalyst, with control of the catalyst acidity, generally measured by the alpha value,
As well as by controlling the reaction conditions, it is possible to vary the properties of the product, in particular the conversion, and thus the amount of the converted fraction from the second stage of the process.

The catalyst used in the hydroisomerization step is a catalyst having a high selectivity for isomerizing waxy linear or nearly linear paraffins to low waxy isoparaffin products. This type of catalyst is bifunctional in nature and comprises a metal component impregnated on a relatively low acid, large pore size porous support. At this stage of the operation, the acidity is kept at a low level to reduce conversion to products having boiling points outside the lubricant boiling range. Generally, prior to metal conversion, the catalyst should have an alpha value of 30 or less, with preferred values being 20 or less (see Example 1).

The alpha value is an approximate indicator of the catalytic cracking activity of a catalyst as compared to a standard catalyst. By Alpha test,
The relative rate constant of the test catalyst (conversion rate of n-hexane per unit volume of catalyst per unit time) is given relative to a standard catalyst having an alpha value of 1 (rate constant = 0.016 sec ^-1 ). The Alpha test is described in U.S. Pat. No. 3,354,078 and the Journal of Catalysis (J.
s), Volume 4, p. 527 (1965); Volume 6, p. 278 (1
966); and Vol. 61, p. 395 (1980), which can be referred to for a description of this test. The experimental conditions of the test used to determine the alpha value referred to herein include variable flow rates and 538 as described in detail in Journal of Catalysis, Vol. 61, p. 395 (1980). Includes constant temperature of ° C.

The hydroisomerization catalyst comprises a large pore zeolite and a metal. Large pore zeolites are supported by a porous binder. Large-pore zeolites usually have at least one open channel consisting of a 12-membered oxygen ring. Large pore zeolites typically have at least one open channel with large dimensions of 7 Å or more. Zeolite beta, Y and mordenite are examples of large pore zeolites.

As disclosed in U.S. Pat.No. 4,419,220, zeolite beta has been found to have excellent activity against paraffin isomerization in the presence of aromatic hydrocarbon compounds, as described above. The preferred hydroisomerization catalyst uses zeolite beta. The low acid form of zeolite beta can also be obtained by synthesizing high siliceous forms of zeolites, for example zeolites having a silica-alumina ratio of 500: 1 or more, and more easily a lower silica-alumina ratio. By steaming the zeolite to the required acidity level. They can also be obtained by extraction with an acid, such as a dicarboxylic acid, as disclosed in US Pat. No. 5,200,168. U.S. Pat. No. 5,164,169 discloses the production of high siliceous zeolite beta using a chelating agent, such as tert-alkenolamine, in a synthesis mixture.

The most preferred zeolites are those which have been severely steamed and have a framework silica-alumina ratio of greater than 200: 1. The silica-alumina ratio is preferably at least 400: 1, and more preferably the silica-alumina ratio is at least 600: 1.

Steaming conditions should be adjusted to achieve the desired alpha value in the final catalyst, and are generally
Use a 100% steam atmosphere at a temperature of 0 ° F to 1100 ° F (427 ° C to 595 ° C). Usually, the steaming is carried out at a temperature of 1000 ° F. (538 ° C.) or higher for 12 to 120 hours, generally 96 hours, to achieve the desired degree of reduction in acidity.

Another method involves replacing a portion of the framework alumina of the zeolite with another trivalent element, such as boron, which results in essentially lower levels of acid activity in the zeolite. Preferred zeolites of this type are those containing the framework boron. Boron is usually added to the zeolite framework before adding the other metals. In this type of zeolite, the framework is mainly
It is made of tetrahedral coordinated silicon, linked by oxygen bridges. A small amount of element (alumina in the case of alumina-silicate zeolite beta) is also coordinated to form part of the skeleton. Zeolite may contain substances in the pores of the structure, but these do not constitute the framework that constitutes the characteristic structure of the zeolite. The term "backbone" boron is defined herein as contributing to the ion exchange capacity of a zeolite from substances that are present in the pores and have little effect on the total ion exchange capacity of the zeolite. It is used to identify substances in the zeolite framework. Zeolite beta is
It has a constraint index in the range of 0.60 to 2.0 at a temperature in the range of 316 ° C to 399 ° C, with a constraint index of 1 or less being preferred.

Methods for preparing high silica zeolite containing framework boron are known and are described, for example, in US Pat. No. 4,269,813. A method for preparing zeolite beta containing skeletal boron is disclosed in US Pat. No. 4,672,049. As described therein, the amount of boron contained in the zeolite can be increased by changing the amount of borate ions in the zeolite-forming solution, e.g. Zeolite Beta can be changed. References can be made to these documents for a description of the method of preparing these zeolites.

The low acid zeolite beta catalyst should contain at least 0.1% by weight of the framework boron, preferably at least 0.5% by weight. Boron can be added to the framework before adding other metals. Typically, the maximum amount of boron is about 5% by weight of the zeolite, and often does not exceed 2% by weight of the zeolite. The skeleton usually contains some alumina. The silica: alumina ratio is usually at least 30: 1 under the conditions of the as-synthesized zeolite. Preferred boron-substituted zeolite beta catalysts are obtained by steaming an initial boron-containing zeolite containing at least 1% by weight of boron (as B ₂ O ₃ ), and the final boron-containing zeolite does not exceed 20 and preferably does not exceed 10. The alpha value is obtained.

Zeolites are complexed with a matrix material to form the final catalyst, for which purpose conventional matrix materials with very low acidity, such as alumina, silica-alumina and silica, are preferred, Alumina such as, for example, alpha boehmite (alpha alumina monohydrate) can also be used, provided that it does not impart a substantial degree of acidic activity to the catalyst formed. Complexing a zeolite with a matrix usually involves
It is carried out at a weight ratio of zeolite: matrix of 80:20 to 20:80, generally 80:20 to 50:50. Conjugation can be accomplished by conventional means, including kneading the materials together and extruding into the desired final catalyst particles. A preferred method for extruding zeolites with silica as a binder is disclosed in U.S. Pat. No. 4,582,815. If the catalyst is steamed to achieve the desired low acidity, the operation is carried out after combining the catalyst with the binder in a conventional manner. A preferred binder for the catalyst to be steamed is alumina.

2. Conditions The conditions used in the hydroisomerization stage depend on the nature of the feed from the first stage and the nature of the product desired from the second stage treatment. If it is desired to maximize the hydroisomerization in the second stage at the expense of hydrocracking,
The temperature needs to be maintained at a relatively low level to minimize the conversion, ie, the conversion to lower boiling products, ie, the bulk conversion. Achieving this requires relatively low temperatures, typically in the range of 200 ° C. (390 ° F.) to 400 ° C. (750 ° F.), but achieves significant overall conversion and hydroisomerization If desired, relatively high temperatures, generally at least 300 ° C. (570 ° F.) and up to 450 ° C. (850 ° F.), will be preferred. The balance of conversion to isomerization is also determined by the general severity of the reaction conditions, including space velocities, which generally range from 0.5 to 10, more generally
It is in the range of 0.5 to 2, typically ¹ hr ^-1 (LHSV).

Because of the low heteroatom content of the resid fraction from the first stage hydrocracking, the process conditions required in the second stage are generally very mild and require relatively low pressure, typically 15000 kPa (2160 psig). Below, in many cases
10000 kPa (1435 psig) or less or 7000 kPa (1000 psig)
It is as follows. This allows the second stage to be converted to existing low pressure equipment, such as catalytic hydrosulfurization (CHD), which is typically operated at pressures below 15,000 kPa (2000 psig).
urization)) within the constraints of the device.

The total pressure of the second stage system can vary up to a maximum value which is basically determined by equipment constraints,
It does not exceed 30000 kPa (4350 psig) and often does not exceed 15000 kPa (2160 psig). At higher hydrogen pressures, the isomerization of the zeolite beta-based catalyst is reduced, because the unsaturated intermediates in the reaction mechanism are more susceptible to shielding by saturated reactants.
Therefore, it is desirable to operate the isomerization reaction at a relatively low hydrogen pressure. In most cases, below 10,000 kPa (1435 psig) or 7000 kPa (1000 n.ll ^-1 (1125-5600sdf / bb
l) Satisfactory results can be obtained in the following, preferably 500 to 1000 n.ll ^-1 (2800 to 5600 sdf / bbl). However, the conditions chosen are, as described above,
It should be selected according to other reaction parameters, including the acidity of the catalyst.

B. Shape Selective Catalytic Dewaxing 1. Catalyst The unconverted resid from the hydrocracking stage contains a greater amount of waxy linear material, n-paraffin, as described above. Non-linear paraffins with higher melting points are also included. These materials may contribute to undesirable pour points in lubricants and the fact that hydrocracker resid materials often have a pour point higher than the target pour point of the product. Need to be separated. To perform this operation without separating the desired isoparaffinic components that contribute to a high viscosity index (VI) in the lubricant product,
A shape-selective dewaxing catalyst is used. This catalyst separates n-paraffins with waxy, somewhat branched paraffins, but leaves more branched chain isoparaffins in the process stream. The shape-selective dewaxing process uses a more selective catalyst for the separation of n-paraffins and paraffins with some branching than isomerization catalysts such as zeolite beta. Shape selective dewaxing is performed as described in US Pat. No. 4,919,788, to which reference can be made for a description of this stage. The catalytic dewaxing step of the process of the present invention is performed using a constrained mesopore size crystalline material, such as a constrained shape selective dewaxing catalyst based on alumina-phosphate. The constrained intermediate pore size crystalline material is:
It has at least one channel of a 10-membered oxygen ring with intersecting channels of an 8-membered ring. Although ZSM-23 is a preferred zeolite for this purpose, other highly shape-selective zeolites, such as ZSM-22, ZSM-48 or synthetic ferrierite ZSM-35, also, especially for lighter materials,
Can be used. Silicoaluminophosphate,
For example, SAPO-11 and SAPO-41 can be used as selective dewaxing catalysts.

Preferred catalysts for use as shape-selective dewaxing catalysts are relatively constrained intermediate pore size zeolites. Such preferred zeolites are described in U.S. Pat.
Measured by the method described in No. 6,218, 1-12
Has a constraint index in the range of These preferred zeolites can also be characterized by sorption properties associated with their relatively constrained diffusion properties. These sorption properties are consistent with zeolites, such as zeolite ZSM-2.
2, ZSM-23, ZSM-35, ferrierite and the like have properties as described in US Pat. No. 4,810,357. These zeolites have pore openings that achieve a particular combination of sorption properties. The combinations are: (1) (sorption, n-hexane at 50 ° C. and o
-For xylene, at a temperature of 80 ° C., measured at a P / P _{0 of} 0.1, based on a volume percentage): the sorption ratio of n-hexane to o-xylene is 3 or more; and (2) At 1000 ° F (538 ° C) and 1 atm, (1000 ° F
With the ratio of the rate constants k _3MP / k _DMB measured at a temperature of about 2 or greater, from a 1/1/1 weight ratio mixture of n-hexane / 3-methylpentane / 2,3-dimethylbutane, 2 It has the ability to selectively degrade 3-methylpentane (3MP) over two-branched 2,3-dimethylbutane (DMB).

The expression “P / P ₀ ” is used, for example, in “The Dynamic Character of Adsorption” by JHdeBoer (second edition,
Oxford University Press (1968)
Year)), and has the same meaning as described in the literature, and is defined as the ratio of the partial pressure of the sorbate at the sorption temperature to the vapor pressure of the sorbate. Target pressure. The ratio of the rate constants, k _3MP / k _DMB, is given by the formula: k = (1 / T _c ) 1n (1 / 1−ε) [where k is the rate constant of each component, and T _c is the contact time And ε is the conversion of each component. ] From the primary velocity in the usual way.

Zeolites that meet these sorption requirements include ferrierite, a natural zeolite, and ZS, a synthetic zeolite.
M-22, ZSM-23 and ZSM-35 are included. These zeolites, when used in the method of the present invention,
It is at least partially in the acid or hydrogen form.

The manufacture and properties of zeolite ZSM-22 are described in U.S. Pat.
No. 4,810,357 (Chester), which may be referred to for such preparations.

Synthetic zeolites ZSM-23 are described in U.S. Patent Nos. 4,076,842 and 4,104,151, to which reference may be made for zeolites, preparation and properties.

Intermediate pore size synthetic crystalline material referred to as ZSM-35 ("Zeolite ZSM-35" or simply "ZSM-35") is described in U.S. Patent No. 4,016,245 for a description of this zeolite and its preparation. Can refer to it. The synthesis of SAPO-11 is described in U.S. Patent Nos. 4,943,424 and 4,440,871. The synthesis of SAPO-41 is
It is described in U.S. Pat. No. 4,440,871. Other examples of intermediate pore size zeolites include ZSM-5 (US Pat. No. 3,702,8
No. 86), ZSM-11 (US Pat. No. 3,709,979), ZSM-2
2, ZSM-23 (U.S. Pat.No. 4,076,842), ZSM-20 (U.S. Pat.No. 3,972,983), ZSM-48 (U.S. Pat.
Nos., ZSM-57 and MCM-22 (US Pat. No. 4,954,325).
No.) is included. The disclosures of these patent documents are incorporated herein by reference. ZSM-5 is included in the HZSM-5 model described above because it can be used without metals.

Ferrierite is a natural mineral, for example, DWBr
eck, ZEOLITE MOLECULAR SIEVES, John Wiley and Sons
(1974), pages 125-127, pages 146, 219, and 625, and this zeolite can be referred to.

The molar ratio of the conjugated metal oxide to silica is from 20: 1 to 200:
One or more suitable molecular sieves can also be used. For example, for HZSM-5, it is advantageous to use a conventional aluminosilicate ZSM-5 having a silica: alumina molar ratio of 25: 1 to 70: 1, but also use values of 70: 1 or more. be able to. Bronstead
sted) A typical zeolite catalyst component having an acid moiety is
In essence, it may comprise crystalline aluminosilicate having a structure of ZSM-5 zeolite with 5 to 95% by weight of silica, clay and / or alumina binder. Other intermediate pore size acidic molecular sieves, such as salicylate, silica aluminophosphate (SAPO) materials,
In particular, it should be understood that intermediate pore size SAPO-11 can also be used as a catalyst.

U.S. Pat. No. 4,908,120 (Bowes et al.) Discloses a catalytic process useful for high paraffin content or high nitrogen level feeds. The method uses a binder-free zeolite dewaxing catalyst, preferably ZSM-5.

Intermediate pore size zeolites are particularly useful in such processes because of their stability under extreme operating conditions, long life and reproducibility. Usually, the zeolite catalyst has a crystal size of from 0.01 μm to 2 μm or more, preferably from 0.02 to 1 μm. ZSM-5 (α value 4
0 or more) can be used for selective decomposition in a metal-free form, but in the case of other intermediate pore size acidic metallosilicates as described above, for use as a hydroisomerization dewaxing catalyst, It must be modified with 0.1-1.0% by weight of noble metal.

ZSM-5 is the only mesopore size zeolite or mesopore size acidic molecular sieve that can actually be used for industrial selective dewaxing without the addition of precious metals. Precious metals
Other intermediate pore size molecular sieves are needed to reduce the aging rate of the catalyst to a practical level. However, the addition of noble metals to ZSM-5 provides hydroisomerizing activity that improves the yield of dewaxed lubricating oil. When noble metals are added to ZSM-23, ZSM-35, SAPO-11 and ZSM-5, the product yield and viscosity index (V
I) for ZSM-23, ZSM-35 and SAPO-11, Z
It has been found that it is generally higher than in the case of SM-5. The choice of which catalyst to use is economic.

The catalyst dimensions can vary widely within the spirit of the invention, depending on the process conditions and the construction of the reactor. The final catalyst preferably has an average maximum dimension of 1 to 5 mm.

The dewaxing catalyst used for shape-selective catalytic dewaxing comprises a metal hydrogenation-dehydrogenation component. Although not necessary to promote the selective cracking reaction, the presence of this component is desirable to promote a particular isomerization reaction that contributes to the synergistic effect of the two-catalyst dewaxing system. It has been found that. The presence of the metal component improves, in particular, the viscosity index (VI) and stability in the product and contributes to slowing the aging of the catalyst. Shape-selective catalytic dewaxing is usually performed in the presence of hydrogen under pressure. The metal is preferably platinum or palladium. The amount of metal component is generally between 0.1 and 10% by weight. Matrix materials and binders can be used as needed.

2. Conditions Shape-selective dewaxing using highly constrained, high-shape-selectivity catalysts is performed by other catalytic dewaxing methods, for example, the same general method as described above for isomerization catalysts. be able to. Thus, conditions are high temperatures and pressures, typically temperatures between 250 ° C and 500 ° C (580 ° F and 930 ° F), more usually temperatures between 300 ° C and 450 ° C (570 ° F and 840 ° F), large For parts, the temperature does not exceed 370 ° C (700 ° F). Pressures up to 3000 psi (20685 kPa), more usually 2
Up to 500psi (17238kPa). Space velocity is 0.1
1010 hr ⁻¹ (LHSV), more usually 0.2-5 hr ⁻¹ . Hydrogen circulation flow rate is 500-1000n.ll ^-1 , more usually 200-400
The range is nll ^-1 . For a broader description of the shape selective catalytic dewaxing process, see US Pat. No. 4,919,788.
No. can be referred to. As already explained, to provide the best temperature control in the reactor,
Hydrogen can be used as an interbed quench. Pt / ZSM-23 is primarily a shape-selective catalyst, but adds incremental isomerization capacity.

The choice of the dewaxing catalyst and the degree of conversion to lower boiling species in the shape-selective dewaxing step depend on the degree of dewaxing desired there, i.e., the unconverted resid material from the hydrocracking step. And the desired pour point. This also depends on the selectivity of the shape-selective catalyst used. At lower product pour points, the use of relatively low selectivity dewaxing catalysts results in higher conversions and correspondingly high hydrogen consumption. Generally, boiling points outside the range of lubricants, e.g.
The conversion to a product with a boiling point of 43 ° C. is at least 5
%, Often at least 10% by weight, and up to 30% by weight conversion is required to achieve only the lowest pour point with the required selectivity of the catalyst. The boiling range conversion on a 650 ° F. + (343 ° C. +) basis is usually in the range of 10-25% by weight.

After the pour point of the product is reduced to a desired value by selective dewaxing, the colored substance is removed and subjected to a treatment such as a hydrogenation treatment to obtain a lubricant product having desired characteristics. Can be. Fractional distillation can also be performed to separate light end materials and to satisfy volatile properties.

The isomerization catalyst, for example, zeolite beta, operates selectively to convert 40% to 90%, more preferably 50% to 80% of the wax in the feed to the isomerization process, and selectively converts bulky wax molecules. It is effective for Shape-selective mesopore size catalysts, in contrast, are more effective at separating the wax in the front end (low boiling component) of the feed, so that the pour point and cloud point of the product, eg, lubricant, Decrease. Intermediate pore size molecular sieves, such as Pt / Z
SM-23 has increased isomerization capacity in addition to shape-selective dewaxing capacity.

FIG. 1 shows one preferred embodiment of the present invention. The feed may be a hydrocarbon feed having a boiling point of 343 ° C. or higher (see description above), before being introduced into the hydrocracking unit 1, from the resid material from the dewaxing step and from the line 2. Mixed with hydrogen. Hydrocracking equipment 1
Has a catalytic hydrocracking zone. This can include a hydrotreating zone alone or a hydrotreating / hydrocracking zone combination. A fixed bed reactor is preferred. The conditions are effective to convert at least 20% of the feed in one pass to a material having a boiling point below the initial boiling point of the feed. The pressure at which the operation of the hydrocracking step of FIG. 1 is performed is substantially higher than the pressure used in the dewaxing step.

The effluent from hydrocracker 1 containing excess hydrogen is:
The hydrocracked product is sent to the high-temperature separation device 3 through the line 4 and separated from hydrogen sulfide and hydrogen contaminated with ammonia. In an alternative, not shown, the effluent from the hydrocracker 1 is cascaded directly to the dewaxing stage without intermediate stage separation. Desulfurization and denitrification often involve the saturation of aromatics occurring in the hydrocracker. The contaminated hydrogen is sent from the hot separator 5 to the cold separator 6, where a substantial portion of the contaminated gas is separated through line 7,
In some cases, it is transferred to offgas. Hydrogen is compressed 8
After being sent to the inlet of the hydrocracker through line 2. The hydrocracked product is stripped of some gaseous components and is transferred to a steam stripper 10 through line 9 and further stripped of gaseous components and light end materials to off-gas through line 11. The hydrocracked product is sent through line 12 to a flash zone of a fractionator 13 where a baffle 14 is provided. The baffle is for separating the hydrocracker bottoms from the dewaxed bottoms. The hydrocracked product is subjected to atmospheric distillation to separate light boiling components such as naphtha and middle distillate products.

The hydrocracker resid material is sent through line 15 to the inlet of an isomerization / dewaxing zone 16 where it is mixed with hydrogen from line 17. The isomerization / dewaxing zone is a large pore zeolite dewaxing catalyst, such as zeolite beta or a shape-selective catalyst (see discussion on dewaxing catalysts above).
It is a fixed bed reactor containing. The effluent from the dewaxing device
As described above, the light gas and the hydrogen are sent to the high-temperature separation device 18 for separation. The gas stream is on line 19
Through the cryogenic separation unit 20. These separators are operated at temperatures ranging from 100 ° F. (38 ° C.) to 200 ° F. (93 ° C.), with contaminated gases being transferred to off-gas through line 21. The contaminated gas from the hydrocracking unit, line 11,
These gases are mixed in line 22. The hydrogen is sent through line 23 to a compression unit 24, where the pressure is increased before returning to the hydrocracker through line 25. The effluent from the hot separator 18 enters the stream stripper 26 via line 27, where gaseous and light components are further separated through line 28 and mixed with the contents of line 22 before entering the flash drum 29, where C ₃
-Some substances having a range of boiling points are separated off-gas by line 30. The remainder of the flash drum effluent is passed through a line 32 to a butanizer 31
Sent to In Butanaiza 31, C ₄ suitable for LPG
The material is separated and the remainder of the flash drum effluent is passed through line 33 to the flash section of fractionator 13 (flash po
rtion). The dewaxed effluent from which the light components have been separated is sent through line 34 to the bottom of fractionator 13 where it is separated from the hydrocracked resid by baffle 14. Materials having boiling points in the naphtha and middle distillate range are separated. The remaining unconverted, dewaxed material is suitable for use as a lubricant base oil material and can be removed for use through line 35 for that purpose. Some or all of the unconverted dewaxed material is recycled by line 36 and mixed with fresh feed before entering the hydrocracker by line 22.

FIG. 2 shows a slightly different embodiment. The configuration of FIG. 2 is aimed at maintaining sufficient separation of all products from the two reaction stages. In some cases, it may be advantageous to carry out the dewaxing / isomerization reaction at elevated pressure.
Instead of a single fractionator, two separate fractionators are
One can be provided downstream of the stripper 2 and the other can be provided downstream of the stripper 6. The stripped hydrocracker effluent enters the fractionator 3 via line 5. The lighter products are separated by atmospheric distillation and the hydrocracker bottoms material is transferred from the bottom of the fractionator through line 15 to the isomerization / dewaxing zone 13. The stripped and dewaxed material is sent from the stripper 6 to the fractionator 4. As shown in FIG. 1, the unconverted resid material can be sent to the lubricant base oil material through line 12 or recycled to the hydrocracker through line 7.

FIG. 3 is an integrated method similar in configuration to FIG. 1 except that a make-up hydrogen compressor is used to supply the one-pass hydrogen stream in the compression stage.

Continued on the front page (72) Inventor Gobel, Kenneth W. No. 5135, Ancinetta Drive, Schnexville, Pennsylvania, United States 18078, USA (72) Inventor Hilbert, Timothy Lee, 1717, Charles Lane, Swell, NJ 80080, New Jersey 72) Inventor Hunter, Michael G. United States 77450 Crystal Bay Drive, 22322, Katie, Texas (72) Inventor Papal, David A United States 08033, 118 Payton Avenue, Haddonfield, NJ 118 (72) Invention Gently, Arthur Earl United States 77079 Tyler Crest 17322 Houston, Texas (72) Inventor Partridge, Randall David United States 08033 Haddonfield, NJ No. 229, Mount Well Avenue (56) Reference JP-A-63-277296 (JP, A) JP-A-59-36194 (JP, A) US Pat. No. 5,275,719 (US, A) (58) Fields investigated ( Int.Cl. ⁷ , DB name) C10G 65/12 C10G 45/58-45/64 C10G 73/38

Claims (10)

(57) [Claims] An integrated hydrotreating process for upgrading gas oil hydrocarbon feedstocks having an initial boiling point of at least 315 ° C into naphtha products, distillate products above the naphtha boiling range and lubricant products. A process comprising: (a) converting a hydrocarbon feedstock to a lower boiling product using an aromatic-selective hydrocracking catalyst having acid functionality and hydrogenation-dehydrogenation functionality. % Under a hydrocracking condition at a conversion of not more than 1% and a hydrogen partial pressure of not more than 10,000 kPa to form a naphtha product, a distillate product and a resid fraction rich in paraffinic components, (b) ) A step of separating a naphtha product, a distillate product and a resid fraction; (c) a step of hydrotreating the separated resid fraction with at least one hydrotreating catalyst for dewaxing; ( A) separating the effluent from the residual oil fraction and a fraction having a boiling point equal to or lower than the residual oil fraction, wherein at least a part of the residual oil fraction is recycled to the step (a); Separating a boiling point fraction into naphtha and distillate products; and (e) using a residual oil fraction not recycled to step (a) as a lubricant raw material. 2. The process according to claim 1, wherein in step (c), the hydrotreating catalyst is dewaxed by hydroisomerization or shape-selective action. 3. The process according to claim 1, wherein in step (c), dewaxing is carried out using a series of hydroisomerization catalysts and a shape-selective catalyst. 4. In step (c), the hydrotreating catalyst is a low acidity large pore zeolite hydroisomerization catalyst, wherein the catalyst is 30
2. The method of claim 1 having an alpha value not exceeding and comprising a metal hydrogenation component. 5. The method of claim 1 wherein the hydrocracking catalyst comprises a dealuminated USY having a silica / alumina ratio of 50: 1 or greater. 6. The method according to claim 4, wherein the hydroisomerization catalyst comprises zeolite beta. 7. The process of claim 1, wherein in step (c), the hydrotreating catalyst comprises a constrained mesoporous crystalline material having a metal hydrogenation-dehydrogenation component for shape selective dewaxing. the method of. 8. The method of claim 7 wherein the constrained mesoporous crystalline material is ZSM-23. 9. The method of claim 1 wherein in step (c), the hydrotreating catalyst comprises HZSM-5 for shape selective dewaxing. 10. In step (b), the effluents of steps (a) and (c) are separated into each product by atmospheric distillation in a single fractionator, and the residual oil from step (b) is formed. A process according to claim 1 wherein the product is kept separate from the resid product of step (d).