KR100339069B1 - Integrated Hydrogenation Treatment with Separate Recirculation Process - Google Patents

Abstract

The present invention relates to the purification of petroleum products, and more particularly to an integrated hydrotreating reaction comprising a hydrocracking step and a subsequent dewaxing step. The boiling material in the middle distillate or lubricating oil range can be dewaxed. Residual oils (and optionally other oils) of each stage are kept separate from one another during processing. Dewaxing can be carried out using a hydroisomerization catalyst or a form-selective catalyst or in turn both. One example uses baffles in the outflow zone of the fractionator to separate the residual oil from one another. As an alternative, the effluent of the hydrocracking step may be treated separately from the effluent of the dewaxing step. The residual oil fraction of the dewaxing stage can be recycled back to the hydrocracking stage for further processing or used as a lubricant base material.

Description

Integrated Hydroprocessing Process with Separate Recirculation Process

Conventional petroleum hydrocracking and catalytic dewaxing are basically described in different process chemistries. Naphthenes and aromatic compounds are easier to hydrocrack than paraffins because of their molecular structure. The energy required to open the ring compound and the saturated aromatic compound is less than the energy needed to break the paraffin bond. As a result, in conventional high conversion recycle hydrocracking reactions, the recycle oil is characterized by a very high pour point which is rich in paraffins and can sometimes exceed 38 ° C. (100 ° F.). A “morph selectivity” catalyst that limits cracking of the ring structure and preferentially converts straight chain paraffins, which results in substantial pour point and molecular weight reduction, can be used for catalytic dewaxing, in particular for the preparation of lubricating oils. Certain types of dewaxing catalysts are designed to convert linear paraffins by isomerization in combination with pyrolysis, resulting in selectivity for intermediate distillates such as kerosene and diesel oil.

The present invention relates to a hydroprocessing integrated reaction using both hydrocracking and catalytic dewaxing. The unconverted hydrocracked bottoms are subjected to dewaxing in the dewaxing stage. The unconverted material is then recycled back to the hydrocracking step in the dewaxing step. Hydrocracked bottoms fraction, the feed to the catalytic dewaxing stage, is generally rich in straight chain paraffins. These straight chain paraffins are isomerized in the dewaxing step to produce side chained paraffins which are suitable gasoline and extra distillate components.

The dewaxing step can utilize a single catalyst or two catalyst systems. Large pore catalysts such as zeolite B can be used to isomerize or pyrolyze n-paraffins to some extent depending on the catalytic activity. Form selectivity, mesoporous catalysts such as ZSM-5 can also be used to pyrolyze n-paraffins.

Hydrotreatment integration reactions using both hydrocracking and catalytic dewaxing steps are well known in the purification art. In US Pat. Nos. 4,851,109 (Chen et al) and 4,764,266 (Chen et al.), High-boiling feeds, such as diesel and catalytically pyrolyzed cycle oils, are hydrocracked in aromatic selective hydrocracking catalysts, resulting in naphtha. And hydrotreating integrated reactions to produce products in the middle distillate range. The unconverted material can be recycled to the hydrocracking stage or passed through a desoldering stage, where large pore catalysts, especially zeolite beta, which have a hydrogenation-dehydrogenation reaction functionality, to produce more kerosene and distillate range products. Is treated as. The recycling process of hydrocracker remnants is used in these patents, but the unconverted material is not recycled to the hydrocracker in the dewaxing stage. In addition, these processes do not illustrate the use of atmospheric distillation columns with baffles in the flash zone to separate residual material from two different process oils, as illustrated in one embodiment of the present invention. .

PCT application WO 95/10578 discloses a process for the preparation of intermediate distillates by dewaxing hydrocarbon-containing materials having boiling points above 343 ° C. (650 ° F.). The feed is hydrocracked as a catalyst containing a large pore zeolite such as Y and a hydrogenation metal component selected from Group VIB or Group VIII of the Periodic Table. The entire effluent of the hydrocracking zone is then deleadized as a catalyst comprising mesoporous zeolite of type ZSM-5 in the presence of hydrogen. In the present invention, only the residual oil of the hydrocracker effluent is subjected to dewaxing.

EP 189648 A1 discloses a process for hydrotreating heavy diesel feed in order to minimize the production of kerosene and distillate. Figure 5 of this patent illustrates hydrocracking of the feed followed by catalytic dewaxing. In this invention, as in the present application, all the effluent of the hydrocracking apparatus is de-leaded, not the residue. In addition, it is not known for recycle as the first hydrocracking step. Two different hydrocracking stages are used.

U.S. Patent No. 5,053,117 (Kyan et al) discloses a process for deleading hydrocracker residue from an MPHC (MPH) process to make naphtha and cases under MDDW conditions or to lubricating oil under less severe conditions of MLDW. Doing. The feed to the dewaxing apparatus can be combined with diesel with high nitrogen content, thereby increasing the octane number of the naphtha product. It is not known about the recycling of dewaxing residues to hydrocrackers. Gasoline octane number can be increased by minimizing olefin production. This can be accomplished by using an MDDW catalyst and minimum hydrogenation reaction conditions.

Several patents illustrate a generalized process diagram for hydrocracking feeds such as heavy diesel oil (or other hydrocarbon feeds with boiling points above 343 ° C. (650 ° F.)), followed by dewaxing to produce lubricants. U. S. Patent No. 4,414, 097 (Chester et al) discloses a process of hydrocracking a hydrocarbon feedstock having a boiling point of at least 343 [deg.] C. (650 [deg.] F.), followed by dewaxing on a catalyst including ZSM-23.

US Pat. Nos. 4,282,271 (Garwood et al) and 4,283,272 (Garwood et al) disclose hydroprocessing of lubricating oils by hydrocracking feedstocks having boiling points above 343 ° C (650 ° F). The hydrocrackate is then subjected to dewaxing. The effluent from the dewaxer is hydrotreated but not recycled to the hydrocracking stage.

FIELD OF THE INVENTION The present invention relates to the purification of petroleum products, and more particularly to boiling in the middle distillate and lubricating oil ranges, which keep the bottoms streams (and optionally other oils) of each stage separately during the treatment process. Integrated hydrotreating reactions of materials (including hydrocracking and subsequent dewaxing with shape selective or hydroisomerization catalysts). The residual oil fraction in the dewaxing stage can be recycled back to the hydrocracking stage for further processing or used as a lubricant base material.

The present invention relates to the purification of petroleum products, and more particularly to an integrated hydroprocessing process, including the hydrocracking step followed by a subsequent step of deleadizing the unconverted hydrocracker residue. Residual oils (and optionally other oils in each stage) are kept separate from each other. Dewaxing can be carried out continuously using hydroisomerization catalysts, or form-selective catalysts or both. One example separates the residual oil from each other using baffles in the outflow zone of the classifier. As an alternative, the effluent from the hydrocracking step can be treated separately from the effluent from the dewaxing step. The residual oil fraction from the dewaxing stage can be recycled back to the hydrocracking stage for further treatment or used as a lubricant base material.

The integrated hydroprocessing process of the present invention can achieve maximum yield at minimal cost by its recycling step. This process allows the dewaxing treatment to be carried out at the same pressure (high pressure dewaxing) or substantially lower pressure as the hydrocracking step, resulting in a conversion with reduced hydrogen consumption. In addition, the product may be taken in a single sorting device or may be taken in one or more sorting devices. In this way, the product obtained by dewaxing can be separated from the product obtained by the hydrocracking step. Although a single fractionation apparatus can be used for normal product recovery, it is also desirable to maintain separation of the unconverted residue product in both reaction stages, since at least a portion of the deleaded effluent can be recycled to the hydrocracker. In one example, the reactor product from the hydrocracking step and the dewaxing step can be removed separately. The combined removal oil is then debutanized. Recover LPG at minimum expense. The reactor liquid product oil is separately replaced by a fractionator (see FIGS. 1 and 3).

In another example, a linked compressor can be used to supply a one-pass hydrogen stream between compression stages. The dewaxing step can be operated at reduced pressure and at a low ratio of oil to gas. The dewaxer outlet hydrogen is then compressed into a hydrocracker stage for make-up, eliminating the need for a separate recycle hydrogen compression plant in the dewaxing stage (FIG. 3).

Dewaxing by isomerization of materials whose boiling points are in the middle midstream range improves the product's room temperature flow characteristics, such as diesel pour point, cloud point, and jet fuel freezing point. In the present invention the paraffinicity of kerosene and diesel products is improved, thereby improving the combustibility of the intermediate distillate.

1 shows the integrated hydrocracking and dewaxing steps of the present invention. Although not shown, the hydrocracking step may also include a dual hydrotreatment hydrocracker catalyst system. Following hydrocracking, the feed is separated into the product of the fractionation apparatus by atmospheric distillation. The residual material of the hydrocracking step is then de-leaded. The fractionation apparatus of FIG. 1 shows a baffle in an outlet area for maintaining the residual oil fraction separately in the hydrocracking step and the dewaxing step. In Fig. 2, the hydrocracked effluent and the dewaxed effluent are treated separately by the use of separate fractionation apparatus and debutulizer. It is preferred if the dewaxing or isomerization will be carried out at high pressure. 1 and 3 show separate removal of reactor product from two steps. The combined removal oil is then debutulized. FIG. 3 shows both hydrocracking and dewaxing steps operated at the same pressure level.

Feed

As feedstocks used in the present invention, other classes of feeds, such as Fischer-Tropsch synthesis or other synthetic processes such as synthetic oil preparation processes such as methanol conversion, are also provided. Although used, it can generally be characterized by a high boiling feed of petroleum. The fraction of fresh raw materials such as shale oil and tar sand can also be treated by the integrated hydroprocessing technology of the present invention. In general, feeds have a large initial boiling point of 345 ° C. (650 ° F.), with relatively high boiling points, typically 205 ° C. (400 ° F.) or higher, such as 230 ° C. (450 ° F.) and in most cases, Will be at least 315 ° C. (600 ° F.). The boiling point properties of the feed, in particular the endpoint, will be determined by the desired product. If a significant amount of lubricating oil is to be produced, the feed itself should contain a significant amount of components above the boiling point of the lubricating oil, typically 345 ° C. (650 ° F.). Thus, when lubricating oil is desired, a significant amount of non-distillable residue will be excluded from the feeds that can be treated, but the feed is generally light oil, i.e. a high boiling distillation feed, typically with an end point typically below 565 ° C (1050 ° F). It will be water. Deasphalted oils (oils obtained from propane deasphalted equipment and suitable for the production of brightstock) can also be used as feeds. Typical feeds that can be treated include light oils such as coke group heavy gas oil, vacuum gas oil, reduced pressure crude oil and atmospheric gas oil. In comparison, if the process is to be used primarily to produce intermediate distillates and naphtha, especially jet fuels, the feed may have a relatively low endpoint since it does not need to preserve higher boiling point components. Thus, for jet fuel production, catalytic pyrolysis cycle oils including light cycle oil (LCO) and heavy cycle oil (HCO), refined slurry oil (CSO) and other catalytic pyrolyzed products are particularly useful for use in the process of the present invention. It is a source. In catalytic pyrolysis processes, cycle oils typically have boiling points of 205-400 ° C. (400-750 ° F.), but light cycle oils may have lower endpoints of 315 ° C. or 345 ° C. (600 ° F. or 650 ° F.). Heavy cycle oils have a higher initial boiling point (IBP), for example 260 ° C. (500 ° F.). The relatively high aromatic content of the cycle oil makes it very suitable for treating the cycle oil in the initial hydrocracking stage of the integrated process sequence of the present invention, in addition the level of reduction currently required for such refractory raw materials is to be treated in the process chart of the present invention. It's made of very attractive stuff.

Because of their high boiling points, the feeds used in the process of the present invention generally contain relatively high proportions of aromatic compounds, in particular polycyclic aromatic compounds, although high boiling paraffins and cycloparaffins are also present in substantial amounts. Catalytically pyrolyzing the feed, the aromatic compound will be substantially dealkylated and will not be further pyrolyzed except in processes that are susceptible to hydrogenation such as the process of the present invention. The process of the present invention is remarkable in its ability to convert refractory aromatic feeds, such as those obtained in catalytic pyrolysis operations, into low flow point, high quality, high paraffinic distillates with low flow point lubricants.

Generally, the aromatic content of the feed used in the process of the present invention will be at least 30% by weight, typically at least 40% by weight and in most cases at least 50% by weight. The balance will be divided into paraffins and naphthenes according to the origin of the feed and its previous treatment. Catalytic pyrolyzed raw materials tend to have higher aromatic content than other feeds and in some cases, the aromatic content may exceed 60% by weight. The feed may be hydrotreated to remove contaminants prior to hydrocracking.

Typical vacuum gas oil (VGO) compositions are shown in Tables 1-4, two hydrotreated (HDT) gas oils are shown in Tables 2 and 3, and also two typical FCC LCO feeds are shown in Tables 5 and 6.

VGO General Characteristics API weight 23.2 Distillation,% by weight 225-345 ° C (440-650 ° F) 7.0 345-400 ° C (650-750 ° F) 17.0 400 ° C + (750 ° F +) 76.0 Sulfur, wt% 2.28 Nitrogen, ppmw. 550 Pour point, ℃ (℉) 18 (95) KV at 100 ° C., cSt. 5.6 P / N / A, wt% 29/21/50

HDT Statfjord VGO Nominal boiling point, ℃ (℉) 345-455 (650-850) API weight 31.0 H, weight% 13.76 S, weight% 0.012 N, ppmw 34 Pour point, ℃ (℉) 32 (90) KV, cSt at 100 ℃ 4.139 P / N / A weight% 30/42/28

HDT VGO Nominal boiling point, ℃ (℉) 345-510 (650-950) API weight 38.2 H, weight% 14.65 S, weight% 0.02 N, ppmw 16 Pour point, ℃ (℉) 38 (100) KV, cSt at 100 ℃ 3.324 P / N / A weight% 66/20/14

Arab hard HVGO API weight 22.2 Hydrogen, wt% 12.07 Sulfur, wt% 2.45 Nitrogen, ppmw 600 CCR, weight percent 0.4 P / N / A, wt% 24 / 25.3 / 50.7 Pour point, ℃ (℉) 40 (105) KV, cSt at 100 ℃ 7.0 Distillation (D-1160),% ℃ ℉ IBP 345 (649) 5 358 (676) 10 367 (693) 50 436 (817) 90 532 (989) 95 552 (1026) FBP 579 (1075)

FCC LCO Characteristics API weight 21.0 TBP, 95%, ℃ (℉) 362 (683) Hydrogen, wt% 10.48 Sulfur, wt% 1.3 Nitrogen, ppmw 320 Pour point, ℃ (℉) -15 (5) KV at 100 ° C - P / N / A, wt% -

FCC LCO Characteristics Distillation,% by weight 215 ° C- (420 ° F) 4.8 215-345 ° C (420-650 ° F) 87.9 345-425 ° C (650-800 ° F) 7.3 425-540 ° C (800-1000 ° F) - 540 ℃ + (1000 ℉ +) - H, weight% 10.64 S, weight% 1.01 N, weight% 0.24 Ni + V, ppmw - CCR, weight percent - HC type, wt% Paraffin Weight% 12.7 Mononaphthene 11.7 Polynaphthene 12.8 Monoaromatic 24.7 Diaromatic 21.7 Polyaromatic 14.3 Aromatic sulfur type 2.1 P / N / A 12.7 / 24.5 / 62.8

CHIBA VGO Feed Characteristics Elemental analysis Boiling Point, ℉ (℃) sulfur 1.55 wt% IBP 569 (298) Nitrogen (total amount) 620 ppm T (10) 667 (353) (Basic) 179 ppm T (50) 812 (433) Hydrogen 12.36 wt% T (90) 1010 (543) Carbon Residue (CCR) 0.82 wt% EP 1021 Residue 8.5 wt% Physical properties Composition, weight% density 0.87 g / cc paraffin 57.6 Pour point 120 ℉ (49 ℃) Aromatic compounds 41.9 API 24.7 Viscosity ＠ 100 ℃ 5.84 cs

Catalyst and Conditions-Initial Hydrocracking Step

The purpose of the hydrocracking carried out in the early stages of the integrated hydrotreating reaction is to provide a feed with a relatively high concentration of paraffins for the treatment of the second stage in the catalyst or form-selective dewaxing catalyst. Thus, the catalyst used in the first stage hydrocracking is a catalyst that is selective for more aromatic compounds than paraffin, although hydrocracking of paraffins is not excluded. Thus, conventional hydrocracking catalysts using a large pore size solid with acidic functionality combined with hydrocracking action are used in this step. The use of large pore size materials is essential for the hydrotreating of high boiling feeds such as diesel, so that bulk, polycyclic aromatic components in the feed can be accessed into the internal pore structure of the catalyst, mainly when characteristic pyrolysis reactions occur. It is thought that. The degree of pyrolysis will occur at the outer surface of the catalyst particles but a major portion of the catalyst surface area is present in the particles. It is essential that the components of the feed to be subjected to pyrolysis approach this internal pore structure. Thus, the acid functionality of the first stage hydrocracking catalyst is determined by a large pore, amorphous material such as alumina, silica-alumina or silica or a large pore size crystalline material, preferably zeolite X, zeolite Y, ZSM-3 , A large pore size aluminosilicate zeolite, such as ZSM-18, ZSM-20, MCM-22 or zeolite beta. Zeolites can be used in a variety of cationic and other forms, preferably with higher stability to resist subsequent loss of degradation and acid functionality under the influence of hydrothermal conditions occurring during hydrocracking. Thus, rare earth exchanged large pore zeolites, for example forms with increased stability, such as REX and REY, besides so-called superstable zeolites Y (USY) and high silica zeolites such as dealalumina Y or dealalumina mordenite are preferred. Do. Particular preference is given to dealumina USY catalysts having a silica / alumina ratio of> 50: 1 in the present invention.

Zeolite ZSM-3 is disclosed in US Pat. No. 3,415,736, Zeolite ZSM-18 is disclosed in US Pat. No. 3,950,496 and Zeolite ZSM-20 is disclosed in US Pat. No. 3,972,983, wherein these zeolites, properties thereof Reference is made to the description of the preparation.

Crystalline components can be incorporated into the amorphous matrix for greater physical strength and the matrix itself can be catalytically active. For example, the matrix can be amorphous alumina or silica-alumina or an acidic clay to participate in the acid-catalyzed pyrolysis reaction that occurs. Typically, the matrix will consist of 40-80% by weight of the catalyst, and more typically 50-75% by weight.

The hydrogenation action can be provided in conventional manner, typically by the use of non- or precious metals in groups VIA and VIII of the periodic table. Such metals include nickel, cobalt, tungsten, molybdenum, palladium or platinum, with nonmetallic nickel, tungsten, cobalt and molybdenum being preferred. Especially as feeds with optionally high levels of contaminants, combinations such as nickel-tungsten, nickel-cobalt and cobalt-molybdenum are particularly useful.

By operating as an aromatic-selective, naphtha directed catalyst under low or moderate harsh conditions, naphtha products with boiling points below 165 ° C. (330 ° F.) can be produced at high rates, which makes it possible to produce high octane gasoline by reforming. It contains a relatively high content of naphthenes. In addition, small to moderate amounts of intermediate distillates are produced as high boiling fractions.

For example, as described in US Pat. No. 4,435,275, if the hydrocracking step is operated in distillate selectivity mode at low or normal hydrogen pressure under low or moderate harsh conditions, the conversion fraction may be combined with a lower proportion of naphtha. Typically will contain a high proportion of middle distillate having a boiling point of 165 to 345 ° C (330 to 650 ° F). Also, as described in US Pat. No. 4,802,402, if a distillate selective catalyst is used, the middle distillate will prevail over the product. The middle distillate fraction can be aromatic in nature if operated at relatively low pressures. In general, the content of aromatic compounds in the product is reduced by using higher pressures. If there is a large amount of aromatics, the middle distillate fraction is generally unsuitable for use directly as a jet fuel or diesel but can be used as a blending component for diesel fuel if combined with other components with greater paraffinicity. It can also be used as fuel oil, either as it is or in combination with other high boiling points. However, these middle distillate fractions are relatively low in sulfur and generally conform to product specifications for use as light fuel oils, for example household heating oils. The unconverted fraction, or the bottom fraction from the first stage hydrocracking at low or normal pressure, is typically waxy and very low in sulfur and nitrogen. These products are valuable and can be marketed as low sulfur light fuel oils, but treatment in the isomerization or dewaxing step significantly increases the distillate yield and quality. The paraffins remaining in the hydrocracker remainder are selectively converted to low boiling products with high content of isoparaffins in the second step.

When operating the hydrocracking step in naphtha selectivity mode, moderate to high harsh conditions will be used. Thus, relatively acidic hydrocracking catalysts or hydroprocessing / hydrocracking catalyst combinations that tend to be naphtha directed will be used at relatively high temperatures, such as 400-450 ° C. (750-850 ° F.) at the top of the hydrocracking range. Temperatures of this magnitude can easily be achieved by controlling hydrogen cooling in the reactor, although severity can also be controlled by appropriate control of the space velocity. In order to promote saturation of aromatics in the feed, high hydrogen partial pressures will be preferred, typically at least 7.000 kPa (1,000 psig) and sometimes at least 10,000 kPa (1,435 psig) will be typical and preferred for naphtha mode operation. . The space velocity (LHSV) will typically be 2 hr ⁻¹ and if a lower operating temperature is desired the value will be 1 hr ⁻¹ . The hydrogen circulation rate will be chosen to maintain the catalytic activity at the selected temperature and is typically at least 350 nll ^-1 (1966 SCF / Bbl) and in most cases 500-1,000 nll ^-1 (2810-5620 SCF / Bbl). The naphtha-selective hydrocracking step is very effective in reducing the sulfur and nitrogen levels in the unconverted resid fraction, which is characterized by low levels of wax and nitrogen as well as high levels of wax paraffins. These paraffins are then selectively isomerized and hydrocracked by a second step operation to produce high quality distillates and kerosene products.

The product distribution change that occurs when naphtha-selective hydrocracking is used in the first stage may be mentioned as follows: The feed containing paraffin, naphthene and aromatics was used to complete the feed with naphtha with significant naphthenic / aromatic properties. Convert to a range of products. The aromatic compound in the feed is reduced by the aromatic selectivity of the first stage catalyst. When the paraffin residue fraction is hydrotreated in the second stage on a zeolite beta based catalyst, paraffins predominantly convert to low flow point isopariffins in the middle distillate boiling range, while a small amount of naphtha is produced. The unconverted fraction is a low flow point product due to isomerization to isoparaffins to its paraffins.

Although the precise conditions to be used in all purification processes will be selected according to the nature of the feed, the operation in naphtha mode is useful by the use of higher acidic catalysts, higher severity and higher hydrogen pressure. A single hydrocracking catalyst or hydrotreating hydrocracking catalyst system can be used. In order to obtain the desired moderate to high harsh conditions suitable for naphtha fabrication, temperature and space velocity (LHSV) interact so that lower temperatures can be used at lower space velocities. Thus, higher temperatures, typically above 315 ° C. (600 ° F.) will be more common, but lower temperatures such as 250 ° C. (480 ° F.) may be used. If a hydrotreating / hydrocracking system is used, the operating temperature will be 650-700 ° F. The upper limit temperature will typically be 450 ° C. (850 ° F.) and temperatures of 400-450 ° C. (750-850 ° F.) provide adequate levels of severity at higher space velocities that result in greater throughput despite accelerated catalyst aging. something to do. The space velocity (LHSV) is typically below 2 hr ⁻¹ and at lower temperatures below 315 ° C. (600 ° F.) and below 1 hr ⁻¹ . Operation in naphtha mode is also useful by using higher hydrogen partial pressures that promote saturation of aromatic compounds and then pyrolyze to ring opening and naphtha. Therefore, a partial pressure of hydrogen of at least 7,000 kPa (1,000 psig), preferably at least 10,000 kPa (1,435 psig) is preferred. Although at the expense of increased hydrogen consumption, higher pressures of 1,160 kPa (2,160 psig) are very useful for saturation of aromatic compounds. The hydrogen circulation rate is typically at least 300 nll ^-1 (at least 1,685 SCF / Bbl), preferably at least 500 nll ^-1 (2,800 SCF / Bbl) and typically 500-1,000 nll ^-1 (2,800-5,600 SCF / BBl )to be. The conversion per pass to the naphtha and low boiling product (165 ° C.) will depend on feed characteristics and conditions, but will generally be 5-25% by weight; The overall bulk conversion per pass (ie conversion to lower boiling product) will generally be 15-80% by weight, usually 20-60% by weight.

It is desirable to maximize the naphthenic / aromatic properties of the naphtha and as a result there may be more room for modification. The residue product from the naphtha-selective hydrocracking step is concentrated in paraffin (relative to the feed). This is because the catalyst used in this step attacks the aromatic compound predominantly over paraffin. The treatment of paraffin-rich residual oil in the subsequent dewaxing step with a hydroisomerization catalyst yields iso-paraffinic liquid products such as jet fuel and low flow distillate in higher overall yields. The residual oil fraction of the hydroisomerization stage is relatively rich in naphthenes and aromatics and is recycled back to the hydrocracker. This step results in naphtha with no naphtha or with higher naphtha yields and greater naphthenic / aromatic properties than the recycling operation of the hydrocracker residue.

The manipulation of hydrocracking in the distillate directed mode generally applies lower severity with the use of less acidic catalysts. Again, temperatures of 250-450 ° C. (480-850 ° F.) may be used although higher space velocities of 2 hr ⁻¹ or less may be suitable to achieve the desired stiffness. If aromatic distillates can be tolerated, low hydrogen partial pressures can be used, such as the hydrocracking process described in US Pat. No. 4,435,275, where the technology for preferred operating conditions in distillate directing mode is indeed true. As described herein, hydrogen pressures of 5,250 to 7,000 kPa (745 to 1,000 psig) range from 250 to 1,000 nll ^-1 (1,400 to 5,600 SCF / Bbl), typically 300 to 800 nll ^-1 (1,685 to 4,500 SCF / It is satisfactory as the hydrogen circulation rate of Bbl). The bulk conversion for 345 ° C .- (650 ° F.) products will generally be 50% or less, typically 30-40%.

The compositional changes resulting from the operation of the hydrocracking step in the low pressure, distillate selection mode may be mentioned as follows: the middle distillate of the feed, including significant boiling paraffin fractions, naphthenes and aromatics, has significant aromatic properties. And a full range of hydrocracking effluent with a relatively small amount of naphtha. The unconverted fraction is again paraffinic because of the aromatic-selective catalyst used in this step. When contacted with the hydrogen isomerization catalyst in the second step, paraffin predominantly attacks paraffins by molecular weight reduction (pyrolysis) to form low-boiling isoparaffins, but slightly changes in boiling point by isomerization to low-boiling isomers immediately below the cut point. There is. In addition, the middle distillate is brought to a low flow point by its isoparaffin content.

As previously presented, the hydrotreatment step precedes hydrocracking to remove contaminants, mainly sulfur, nitrogen and metals present, thus creating a dual hydrotreatment / hydrocracking system. Hydrotreating conditions and catalysts will be selected as conventional for this purpose. Interstage separation of inorganic sulfur and nitrogen may be performed or omitted, such as 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 residues are transferred to the dewaxing or isomerization step. From this step the bottoms can be recycled to a hydrocracker to produce naphthas rich in aromatics and naphthenes.

When distillate selective hydrocracking is used in the hydrocracking stage, the converted fraction will have a relatively high boiling point of paraffin, i.e. 225-345 ° C. (440-650 ° F.), at the end of the end. Will contain oil. If the cut points for separating the part fed to the second step are adjusted accordingly, these paraffins will be transferred to the lower boiling range by isomerization and hydrocracking under the condition that the hydroisomerization catalyst is used in the dewaxing step. Therefore, it may be desirable to supply a few conversion fractions in addition to the unconverted fractions in the second step. The cut point may be as low as 200 ° C. (390 ° F.) but will generally be at least 225 ° C. (440 ° F.) and will most likely be at least 315 ° C. (600 ° F.), generally 345 ° C. (650 ° F.).

Dewaxing Step

Dewaxing can occur using a hydroisomerization process, a form-selective dewaxing process, or both in turn. Both hydroisomerization and form-selective dewaxing are described below.

A. Conditions and Catalysts Using Hydroisomerization

1. Catalyst

This step is necessarily a hydrocracking or hydroisomerization step using a catalyst that combines acidic and hydrogenation functionality based on large pore zeolites such as zeolite beta. Hydrogen functionality may be provided by base metals or precious metals. Nickel, tungsten, cobalt, molybdenum, palladium, platinum or combinations of these metals are suitable. Particular preference is given to nickel-tungsten combinations with fluoride. Acidic functionality is preferably provided by the zeolite beta described in U.S. Patent No. RE 28,341, which is a known zeolite and which is incorporated by reference for a description of the zeolite, its preparation and properties. Hydrotreatment catalysts based on zeolite beta are described in the U.S. Pat. Application 379,421 and its European counterpart EU 94827. As described in these patents, the preferred form of zeolite beta for use in hydrotreating is the high silica form, with a silica-alumina ratio of at least 30: 1 (structural). In exchange for pyrolysis, a silica: alumina ratio of at least 50: 1 and preferably at least 100: 1 or 100: 1 or more, such as 250: 1, 500: 1, can be used to maximize the paraffin isomerization reaction. have. Thus, the use of a suitable silica: alumina ratio of the catalyst is typically used in the properties of the product, in particular in the conversion rate and therefore in the process of step 2, such as the control of catalytic acidity, as measured by the control of the alpha value and the reaction conditions. It can be used to change the amount of oil converted.

The catalyst used in the hydroisomerization step is a high selectivity for isomerizing the linear or near linear paraffins in the wax phase into the isoparaffin product with less wax phase. This type of catalyst is bifunctional in nature and comprises a metal component on a large pore size porous carrier of relatively low acidity. The acidity is maintained at a low level in order to reduce the conversion to boiling products outside the boiling point of the lubricating oil during this operating step. In general, the catalyst should have an alpha value of 30 or less before metal addition, with a preferred value of 20 or less (reference example 1).

The alpha value is a rough indication of the catalytic pyrolysis activity of the catalyst compared to the standard catalyst. The alpha test provides the relative rate constant of the test catalyst (straight chain hexane conversion per volume of catalyst per unit time) for the standard catalyst with alpha as 1 (rate constant = 0.016 sec ⁻¹ ). Alpha tests are described in US Pat. No. 3,543,078 and J. Catalysis, 4, 527 (1965), to which reference is made to the description of this test; 6, 278 (1966). The experimental conditions of the test used to determine the relevant alpha value herein include a constant temperature of 538 ° C. and a variable flow rate as described in detail in J. Catalysis, 61, 395 (1980).

Hydroisomerization catalysts include large pore zeolites and metals. Large pore zeolites are supported by a porous binder. Large pore zeolites typically have at least one pore channel composed of 12 membered oxygen rings. Large pore zeolites typically have at least one pore channel having a juke size of at least 7 mm 3. Zeolite beta, Y and mordenite are examples of large pore zeolites.

As indicated above, zeolite beta is used as the preferred hydroisomerization catalyst because it is known that zeolite has significant activity in paraffin isomerization in the presence of aromatic compounds, as described in US Pat. No. 4,419,220. . The low acidity form of zeolite beta can be obtained, for example, by synthesis of high silicate forms of zeolites with a silica-alumina ratio of at least 500: 1, or more easily, evaporate low rates of silica-alumina zeolite to the desired acidity level. This can be achieved by steaming. They can also be obtained by extraction with an acid such as dicarboxylic acid, as described in US Pat. No. 5,200,168. U. S. Patent No. 5,164, 169 describes the preparation of high silicate zeolite beta using chelating agents such as tertiary alkenolamines in synthetic mixtures.

Most preferred zeolites are those that are heavily evaporated and have a silica-alumina ratio of at least 200: 1. Preferably the silica-alumina ratio is at least 400: 1 and more preferably the silica-alumina ratio is at least 600: 1.

Evaporation conditions must be adjusted to obtain the desired alpha value in the final catalyst and typically employ a 100% steam atmosphere at temperatures of 800-1100 ° F. (427-595 ° C.). Typically, evaporation will be performed at temperatures of 1000 ° F. (538 ° C.), 12-120 hours, typically 96 hours, to achieve the desired acidity decrease.

Another method is to replace the aluminum framework of the zeolite with another trivalent element, such as boron, which results in lower native acid activity levels in the zeolite. Preferred zeolites of this type are zeolites containing a boron skeleton. Boron is typically added to the zeolite skeleton before adding other metals. In this type of zeolite, the backbone consists mainly of silicon tetrahedra coordinating with and interconnecting oxygen bridges. Small amounts of elements (alumina in the case of alumina-silicate zeolite beta) are also coordinated to form part of the backbone. Zeolites also contain materials that are within the pores of the structure, although the pores of the structure do not form part of the backbone that constitutes the zeolite's characteristic structure. The "skeleton" boron is used to distinguish between materials in the backbone of the zeolite that are being demonstrated by contributing ion exchange capacity to the zeolite from materials present in the pores and having no effect on the overall ion exchange capacity of the zeolite. Zeolite beta preferably has a constraint index of 1 or less, but has a constraint index of 0.60 to 2.0 at a temperature of 316 ° C to 399 ° C.

Processes for preparing high silica content zeolites containing skeletal boron are known and are described, for example, in US Pat. No. 4,269,813. As is known herein, the amount of boron contained in the zeolite can be varied by incorporating different amounts of borate ions in the zeolite-forming solution, for example, by varying the amount of boric acid relative to the forces of silica and alumina. Reference is made to these documents for a description of how these zeolites can be prepared.

The low acidity zeolite beta catalyst contains at least 0.1% by weight of skeletal boron, preferably at least 0.5% by weight of boron. Boron may be added to the backbone before adding other metals. Typically, the maximum amount of boron will be about 5% by weight of the zeolite and in most cases is no greater than 2% by weight of the zeolite. The backbone will typically include some alumina. The silica: alumina ratio will typically be at least 30: 1 under the conditions of the synthesized zeolite. Preferred boron-substituted zeolite beta catalysts are prepared by evaporating the first boron-containing zeolite containing at least 1% by weight of boron (as B ₂ O ₃ ) so that the final alpha value is not greater than 20 and preferably not greater than 10. .

Zeolites are complexed with matrix materials to form a finished catalyst and for this purpose, although aluminas such as alpha boehmite (alpha alumina monohydrate) can also be used, conventional such as alumina, silica-alumina and silica Very low acidity matrix materials of are suitable, provided they do not impart substantial acidic activity to the matrixing catalyst. Zeolites are complexes with the matrix in an amount of from 80:20 to 20:80, typically 80:20 to 50:50, as zeolite: matrix. Compounding can be performed by conventional means, including mixing the materials together well and then extruding the desired final catalyst particles. Preferred methods for extruding silica and zeolite as binders are described in US Pat. No. 4,582,815. If the catalyst is to be evaporated in order to obtain the desired low acidity, it is carried out after combining the catalyst with the binder, as is conventional. Preferred binders for the evaporated catalyst are alumina.

2. Conditions

The conditions used in the hydroisomerization step will depend on the feed from the first step and the properties of the desired product from the second step treatment. In order to maximize hydroisomerization in the second stage at the expense of hydrocracking, the temperature must be kept at a relatively low level to minimize conversion to lower boiling products, ie bulk conversion. To achieve this, relatively low temperatures are required, typically from 200 ° C. (390 ° F.) to 400 ° C. (750 ° F.), while relatively high temperatures, typically at least, are necessary to achieve significant bulk conversion and hydroisomerization reactions. Preference is given to 300 ° C. (570 ° F.) to 450 ° C. (850 ° F.) or less. The balance of conversion compared to the isomerization reaction is also measured by the general severity of the reaction conditions, including space velocity (LHSV), which is typically 0.5 to 10 and more typically 0.5 to 2, typically 1 hr ⁻¹ . Will be.

Since the residual oil fraction from the first stage hydrocracking is low in heteroatoms, the desired process conditions in the second stage are generally very mild and relatively low pressure, typically below 15,000 kPa (2160 psig) and in most cases 10,000 kPa (1,435 psig). ) Or 7,000 kPa (1,000 psig) or less. This makes it possible to perform the second step within the limits of existing low pressure devices, for example catalytic hydrodesulfurization (CHD) devices, which typically operate at pressures up to 2000 psig.

In the second stage, the overall system pressure can vary below the maximum measured primarily by plant limits and, in general, for this reason will not exceed 30,000 kPa (4350 psig) and in most cases will not exceed 15,000 kPa (2160 psig). Will not. Isomerization of zeolite beta based catalysts is reduced at higher hydrogen pressures since the unsaturated intermediates in the reaction mechanism are further intercepted by the saturation reaction. Therefore, isomerization is preferred at relatively low hydrogen pressure. In many cases, satisfactory results can be obtained below 10,000 kPa (1,435 psig) or even below 7,000 kPa (1,000 psig). Hydrogen circulation rates of 200 to 1,000 nll ⁻¹ (1,125 to 5,600 SCF / Bbl), preferably 500 to 1,000 nll ⁻¹ (2,800 to 5,600 SCF / Bbl), are suitable. However, the selection conditions should be selected according to other reaction variables, including the acidity of the catalyst, as described above.

B. Form-Selective Catalytic Dewaxing

1. Catalyst

The unconverted bottom oil from the hydrocracking step described previously contains more waxy straight chain, n-paraffins. There are also paraffins that are not straight chains of higher melting point. Because they give a pour point that is not useful for lubricating oil, and because the hydrocracker residue has a pour point above the target pour point for the product, it is necessary to remove these waxy components. In order to remove the waxy component without removing the desired isoparaffin component that contributes to the high V.I. of the lubricating oil component, a form-selective dewaxing catalyst is used. This catalyst removes n-paraffins, like waxy, slightly branched paraffins, while leaving more sidechain isoparaffins in the process oil. The form-selective catalyst dewaxing process uses an isomerization catalyst, such as a catalyst that is more selective for the removal of n-paraffins and slightly branched paraffins than zeolite beta. Form selective dewaxing is performed as described in US Pat. No. 4,919,788, which is incorporated by reference for the description of this step. The catalyst dewaxing step in this process is carried out as a form-selective dewaxing catalyst constrained based on constrained mesoporous crystalline material, such as alumina-phosphate. The constrained intermediate crystalline material has at least one channel of a 10-membered oxygen ring with a blocking channel having an 8-membered ring. Other solid-selective zeolites such as ZSM-22, ZSM-48 or synthetic ferrierite ZSM-35 can also be used as especially harder raw materials, but ZSM-23 is the preferred zeolite for this purpose. Silicoaluminophosphates such as SAPO-11 and SAPO-41 can be used as the selective dewaxing catalyst.

Preferred catalysts for use as form-selective dewaxing catalysts are relatively confined mesoporous zeolites. Such preferred zeolites have a constraint index of 1-12, as measured by the method described in US Pat. No. 4,016,218. These preferred zeolites are also characterized by specific sorption properties with respect to their relatively constrained diffusion properties. These sorption characteristics are those shown in US Pat. No. 4,810,357 for zeolites such as zeolites ZSM-22, ZSM-23, ZSM-35 and ferrierite. These zeolites have a specific combination of sorption characteristics, i.e., n for o-xylene, measured at a temperature of 50 ° C. for n-hexane and 80 ° C. for o-xylene as (1)% by volume. -1 of n-hexane / 3-methyl-pentane / 2,3-dimethylbutane, with the sorption ratio of hexane exceeding 3 and (2) the rate constant ratio k _3MP / k _DMB measured at a temperature of 1000 ° F. With pore openings from the / 1/1 weight ratio mixture to obtain the ability to selectively pyrolyze 3-methylpentane (3MP) over 2,3-dimethylbutane (DMB) with double side chains at 1000 ° F and atmospheric pressure have.

The indication "P / Po" is consistent with the conventional meaning described in the literature, for example in "The Dynamical Character of Adsorption" by JH deBoer, 2nd Edition, Oxford University Press (1968) and for the vapor pressure of sorbate at sorption temperature. Relative pressure, defined as the partial pressure ratio of sorbate. The rate _constant , k _3MP / k _DMB, is determined from the first order reaction rate by the following equation in the usual manner:

k = (1 / T _c ) ln (1 / 1-ε)

Where k is the rate constant for each component, T _c is the contact time and ε is the fractional conversion of each component.

Zeolites matching these sorption conditions include natural zeolite ferrierite and synthetic zeolites ZSM-22, ZSM-23 and ZSM-35. These zeolites are at least partially in acid or hydrogen form when they are used in the present process.

The preparation and properties of zeolite ZSM-22 are described in US Pat. No. 4,810,357 (Chester), to which reference is made.

Synthetic zeolite ZSM-23 is described in US Pat. Nos. 4,076,842 and 4,104,151, which are incorporated by reference for descriptions of such zeolites, their preparation and properties.

Mesoporous-sized synthetic crystalline materials, referred to as ZSM-35 ("zeolite ZSM-35" or simply "ZSM-35"), are described in US Pat. No. 4,016,245, which is incorporated by reference for the description of such zeolites and their preparation. It is. The synthesis of SAPO-11 is described in US Pat. Nos. 4,943,424 and 4,440,871. The synthesis of SAPO-41 is described in US Pat. No. 4,440,871. Other examples of mesoporous zeolites are 5 (US Pat. No. 3,702,886), ZSM-11 (US Pat. No. 3,709,979), ZSM-22, ZSM-23 (US Pat. No. 4,076,842), ZSM-20 (US Pat. No. 3,972,983). ), ZSM-48 (US Pat. No. 4,375,573), ZSM-57, and MCM-22 (US Pat. No. 4,954,325). The specification of these patents belongs to the reference in the present invention. ZSM-5 is included as it can be used in HZSM-5 form as described above without metal.

Ferrierite is described in the literature, for example in D.W. Natural minerals, as described in Breck, ZEOLITE MOLECULAR SIEVES, John Wiley and Sons (1974), pages 125-17, 146, 219 and 625.

Suitable molecular sieves may be used having a molar ratio of coordinating metal oxide to silica of 20: 1 to 200: 1, or more. For example, as HZSM-5, a ratio of 70: 1 or more can be used, but it is useful to use conventional aluminosilicate ZSM-5 with a silica: alumina molar ratio of 25: 1 to 70: 1. Typical zeolite catalyst components with Bronsted acid sites may consist essentially of crystalline aluminosilicates having a ZSM-5 zeolite structure, with 5 to 95% by weight of silica, clay and / or alumina binders. Other mesoporous acidic molecular sieves such as salicylate, silica-aluminophosphate (SAPO) materials, in particular mesoporous SAPO-11, can be used as catalyst.

U.S. Patent 4,908,120 (Bowes et al) discloses a catalytic process useful for feeds with high levels of paraffin or high levels of nitrogen. This process uses a binder free zeolite dewaxing catalyst, preferably ZSM-5.

Mesoporous zeolites are particularly useful in the process because of their regeneration, long life and stability under extreme conditions of operation. Typically zeolite crystals have a crystal size of at least 0.01 to 2 microns, with 0.02-1 microns being preferred. ZSM-5 (≧ 40 alpha) may be used in metal free form for selective pyrolysis, but in the case of the other mesoporous acidic metal-silicates described above, these may be used as hydroisomerized dewaxing catalysts in the range of 0.1 to 1.0 precious metals. It is necessary to modify the weight percentage.

ZSM-5 is the only mesoporous zeolite or mesoporous acid molecular sieve that is practical for use in commercial selective dewaxing without the addition of precious metals. Precious metals are needed as other mesoporous molecular sieves to reduce the catalyst aging rate to substantial levels. However, addition of noble metals to ZSM-5 provides hydroisomerization activity which increases the yield of de-leaded lubricating oil. When adding precious metals to ZSM-23, ZSM-35, SAPO-11, and ZSM-5, product yields and VIs are generally found to be greater for ZSM-23, ZSM-35, and SAPO-11 than for ZSM-5. Paid. The choice of which catalyst to use is one of the economic aspects.

The catalyst size can vary greatly within the inventive concept depending upon the process conditions and reactor structure. Preferred catalysts with an average maximum size of 1 to 5 mm are preferred.

The dewaxing catalyst used in the form-selective catalyst dewaxing comprises a metal hydrogenation-dehydrogenation component. While it may not be necessary to catalyze the selective pyrolysis reaction entirely, it has been found that the presence of this component promotes a constant isomerization reaction that contributes to the synergistic effect of the two catalyst dewaxing systems. The presence of the metal component leads to product improvement, in particular VI, and stability as well as to the delay of catalyst aging. Form-selective, catalytic dewaxing is typically carried out in the presence of hydrogen under pressure. The metal will preferably be platinum or palladium. The amount of metal component will typically be from 0.1 to 10% by weight. Matrix materials and binders may be used as needed.

2. Conditions

The highly constrained, form-selective dewaxing with solid-selective catalyst can be carried out in the same general manner as those described above for other catalyst dewaxing processes, such as those using isomerization catalysts. Thus conditions will consist of elevated temperature and pressure with hydrogen, typically at temperatures of 250 to 500 ° C. (580 ° to 930 ° F.), more typically 300 to 450 ° C. (570 ° to 840 ° F.) and mostly Is not greater than 370 ° C. (700 ° F.). The pressure extends below 3,000 psi (20,685 kPa), more typically 2500 psi (17,238 kPa). The space velocity expands to 0.1 to 10 hr ⁻¹ (LHSV), more typically 0.2 to 5 hr ⁻¹ . The hydrogen circulation rate is from 500 to 1000 nll ^-1 , more typically from 200 to 400 nll ^-1 . For a more extensive description of the form-selective catalyst dewaxing step, see US Pat. No. 4,919,788. As previously presented, hydrogen can be used for interbed cooling to maximize temperature control in the reactor. Pt / ZSM-23 is primarily a form-selective catalyst, but adds an increase in isomerization capacity.

The degree of conversion to the low boiling point type and the choice of dewaxing catalyst in the form-selective dewaxing step will depend in this respect on the desired dewaxing range, ie the difference between the target pour point and the pour point of the unconverted resid from the hydrocracking step. It will also depend on the selectivity of the form-selective catalyst used. At lower product pour points and with relatively low selectivity dewaxing catalysts, higher conversions and correspondingly higher hydrogen consumption will occur. Generally outside the lubricating oil range, the conversion to boiling products, for example at 315 ° C.—more typically 343 ° C.—will be at least 5% by weight, in most cases at least 10% by weight, and conversions of up to 30% by weight It is only necessary to obtain the lowest pour point as a catalyst of the desired selectivity. Conversion rates in the boiling range at the 650 ° F. (343 ° C. +) basis will typically be 10-25% by weight.

After the pour point of the product has been reduced to the desired value by selective dewaxing, treatment such as hydrotreating may be performed to remove the pigment and produce a lubricating oil product of the desired properties. Sorting can be used to remove light ends and conform to volatile specifications.

Isomerization catalysts, such as zeolite beta, are effective at selectively converting bulk wax molecules when manipulated to convert 40% to 90%, more preferably 50% to 80%, of the wax in the feed to the isomerization process. In comparison, the form-selective mesoporous size catalyst is more effective at removing wax from the front end material (low boiling point component) of the feed, thereby reducing the clouding and pour points of products such as lubricants. Medium pore size molecular sieves such as Pt / ZSM-23 have an incremental isomerization capability in addition to the form-selective dewaxing capability.

1 illustrates one preferred embodiment of the present invention. A feed (see description above), which may be a hydrocarbon raw material having a boiling point of 343 ° C. or more (see description above), is combined with hydrogen from line 2 before being introduced into the retentate material and hydrocracker 1. The hydrocracker 1 comprises a catalytic hydrocracking zone. It may also include hydrotreating alone or hydrotreating / hydrocracking combination zones. Fixed bed reactors are preferred. The conditions are effective to convert at least 20% of the feed in a single pass to a material having a boiling point below the initial boiling point of the feed. The pressure at which the hydrocracking step of FIG. 1 is operated is substantially greater than the pressure used for the dewaxing step.

The effluent containing excess hydrogen from the hydrocracker (1) is passed to the hot separator (3) via line (4), where the hydrocrackate is separated from hydrogen contaminated with hydrogen sulfide and ammonia. In a separate example, not shown, the effluent from the hydrocracker 1 is branched directly to the dewaxing stage without any intermediate separation. Desulfurization and denitrification reactions frequently involve saturation of aromatic compounds generated in hydrocrackers. The contaminated hydrogen is passed from the hot separator 3 to the cold separator 6 and a substantial portion of the contaminated gas is removed through the line 7 and eventually passed off gas. Hydrogen is passed through a compressor (8) and then returned via line (2) to the hydrocrack inlet. The hydrocrackate is passed through line 9 to a steam stripper 10 with some gas removed for further removal of the gas and light end, and the gas and light end end to line 11 Pass through gas through. The hydrocrackate is passed through line 12 to the outlet area of fractionation device 13 in which baffle 14 is installed. The baffle aims to separate the hydrocracker bottoms from the dewaxer bottoms. Hydrocracking is distilled under atmospheric pressure to remove lighter boiling components such as naphtha and middle distillates.

The hydrocracker bottoms are passed through line 15 to the inlet of the isomerization / dewaxing zone 16 where they combine with hydrogen from line 17. The isomerization / dewaxing zone is a fixed bed reactor containing large pore zeolite dewaxing catalysts or form-selective catalysts (see above for dewaxing catalysts) such as zeolite beta. As previously described, the dewaxer effluent is passed through hot separator 18 for removal of hydrogen and light gases. Gas-containing oil is passed through the line 19 to the cold separator 20. These separators are operated at temperatures between 100 ° F. (38 ° C.) and 200 ° F. (93 ° C.), where contaminating gases are passed through line 21 to off gas. Hydrogen is passed through the line 23 to the compressor 24 where the pressure is raised before returning to the hydrocracker via line 25. The effluent from the hot separator 18 enters the vapor stripper 26 through line 27 for further removal of the gas and light end through the line 28 and of the line 22 prior to entering the outlet drum 29. Combine with the gas and the hard end, where some of the material boiling in the C ₃ -range is removed via line 3 with off gas. The effluent liquid of the remaining effluent drum is passed to the butaneizer 31 by the line 32. In the butcher 31, C ₄ -material, suitable for LPG, is removed and the effluent of the remaining effluent drum is passed by way of line 33 to the outlet of the fractionation device 13. Via line 34, the deleaded effluent, along with the removed hard end, is transferred to the bottom of the fractionation apparatus 13 for atmospheric distillation and separated from the hydrocrackate residue by baffles 14. Remove boiling material from naphtha and middle distillate ranges. Residual unconverted dewaxed material is suitable for use as a lubricant base material and can be removed by line 35 for this purpose. Some or all of the unconverted dewaxing material may be recycled by line 36 and combined with fresh feed prior to entering the hydrocracker by line 25.

2 provides a slightly different example. The design of FIG. 2 aims to maintain complete separation of all products from the two reaction steps. In some cases it may be beneficial to carry out the dewaxing / isomerization reaction at high pressure. Instead of a single sorting device, two separate sorting devices are used, one to the stripper 2 and the other to the stripper 6. The stripped hydrocracker effluent is introduced via line 5 into fractionation apparatus 3. Atmospheric distillation removes the lighter product and transfers the hydrocracker bottoms through the line 15 to the isomerization / decapsulator zone 13 at the bottom of the fractionation unit. The stripped and dewaxed material is passed from stripper 6 to fractionation device 4. As shown in FIG. 1, the unconverted residual oil material can be passed through line 12 to a lubricant based raw material or recycled through line 7 to a hydrocracker.

FIG. 3 is an integrated process diagram similar to that of FIG. 1 except for using an assembled hydrogen compressor to provide a one-pass hydrogen stream between compression stages.

Claims (10)

Integrated hydroprocessing method for quality improvement of diesel hydrocarbon raw materials with initial boiling point of 315 ° C. or higher to naphtha products, distillates and lubricating oil products boiling above the naphtha range, including the following steps: (a) a naphtha product by hydrocracking a hydrocarbon feedstock in an aromatic-selective hydrocracking catalyst with acid functionality and hydrogenation-dehydrogenation functionality under hydrocracking conditions of less than 10,000 kPa of hydrogen partial pressure and less than 50% by volume conversion to a low boiling point product, Forms a distillate, and a residual oil fraction rich in paraffin components; (b) separating the naphtha product, distillate and bottoms fraction; (c) hydrotreating the separated residual oil fraction in one or more hydroprocessing catalysts for dewaxing; (d) separating the effluent of step (c) into an oil fraction having a residual oil fraction and boiling point below the residual oil range, wherein part or all of the residual oil fraction is recycled to step (a) and the low boiling fraction is separated into naphtha and distillate; (e) Residual oil fraction not recycled to step (a) is used as lubricating oil basestock. The process according to claim 1, wherein in step (c) the hydrotreating catalyst is subjected to dewaxing by hydroisomerization or by form-selective means. 2. The process according to claim 1, wherein in step (c), the dewaxing is in turn carried out on a hydroisomerization catalyst and a form-selective catalyst. The process according to claim 1, wherein in step (c), the hydrotreatment catalyst is a low acidity, large pore zeolite hydroisomerization catalyst, the catalyst containing a metal hydrogenation component with an alpha value of no greater than 30. The process of claim 1 wherein the hydrocracking catalyst comprises dealuminated USY having a silica / alumina ratio> 50: 1. The process of claim 4 wherein the hydroisomerization catalyst comprises zeolite beta. The process of claim 1, wherein in step (c) the hydrotreating catalyst comprises a confined mesoporous crystalline material having a metal hydrogenation-dehydrogenation component for form-selective dewaxing. 8. The method of claim 7, wherein the confined mesoporous crystalline material is ZSM-23. The process of claim 1, wherein in step (c), the hydrotreating catalyst comprises HZSM-5 for form-selective dewaxing. The process of claim 1, wherein in step (b), the effluent of steps (a) and (c) is separated into the product by atmospheric distillation in a single fractionation apparatus, in which the residue product of step (b) is subjected to step (d). Keeping separate from the residue product).