Wax oil hydrotreating method

Abstract

The invention discloses a wax oil hydrotreating method. Reaction products are timely removed and separated, and catalyst grading setting is performed according to different components, so that raw oil can be subjected to echelon targeted hydrogenation reaction in the reactor, and the problem that the blending ratio is limited due to the fact that the hydrogenation depth cannot be better controlled during blending of catalytic diesel oil and wax oil raw materials is solved; the aromatic hydrocarbon content in heavy naphtha in a hydrocracking product is greatly increased, and the product quality of diesel oil and tail oil is improved.

Description

Technical Field

[0001] The invention belongs to the field of petroleum refining, and in particular relates to a wax oil hydrogenation treatment method. Background Art

[0002] With the adjustment of domestic economic structure and the slowdown of economic growth, the growth rate of diesel consumption has slowed down, and the market demand for jet fuel and chemical raw materials has gradually recovered. In particular, the energy structure transformation period is coming, and the change in consumption structure will bring challenges to the adaptability of China's refining unit structure. Hydrocracking technology is the core of the combination of "oil, chemical fiber" in the refining structure, which can produce high-quality jet fuel, diesel, lubricating oil base oil raw materials, catalytic reforming raw materials and ethylene raw materials. Wax oil hydrogenation units can also blend catalytic cracking diesel, coking diesel and other inferior diesel to produce chemical materials, high-quality gasoline and other products, realizing the high added value of surplus diesel.

[0003] For example, patents CN103773495A, CN101875856A, CN102465035A and CN1896192A disclose the process technology of blending catalytic cracking light cycle oil in the wax oil hydroprocessing process, the main purpose of which is to produce high-quality catalytic cracking feedstock, or to achieve clean production of the catalytic cracking unit by coupling the catalytic cracking light cycle oil to circulate between the wax oil hydroprocessing unit and the catalytic cracking unit.

[0004] CN112410068A discloses a wax oil hydrocracking device and a method for producing 5# industrial white oil. The mixed raw materials composed of catalytic diesel and coking diesel mixed with atmospheric and vacuum light wax oil are sequentially fed into a hydrorefining reactor and a hydrocracking reactor; the cracking reaction products are subjected to gas-liquid separation in a cold and hot high and low separator; the reaction generated oil enters a hydrogen sulfide stripping tower and then enters a flash tank; the reaction generated oil at the bottom of the flash tank is heated and fed into a fractionating tower for cutting to obtain naphtha, jet fuel, industrial white oil and tail oil respectively.

[0005] In the above patent, since the wax oil is directly hydrogenated after being mixed with the refined oil, the hydrogenation depth of the refined oil cannot be effectively controlled, the blending ratio is limited, and the flexibility of the device is affected.

[0006] CN109988630A discloses a wax oil hydrogenation method and system, which hydrogenates the logistics in steps, performs gas-liquid separation between the two steps of hydrogenation, extracts the hydrogenated gas phase logistics of the wax oil raw material after hydrogenation, and sends it and catalytic cracking light cycle oil to a separate hydrofining reaction, so as to control the hydrogenation depth of the catalytic cracking light cycle oil, and finally increase the content of aromatics in the catalytic cracking gasoline, thereby increasing the octane number of the catalytic cracking gasoline; in this method, although the hydrogenation depth of the catalytic diesel is controlled by hydrogen sulfide and ammonia in the oil generated by hydrogenation of the wax oil, due to the plug flow reaction process of the fixed bed reactor, there are still many fractions that are over-ring-opened or cracked, which limits the further increase of the aromatic content in the gasoline.

[0007] Patent CN110938466A controls the cracking depth of different distillate oils by using different pore sizes, different molecular sieves and different active metal contents in the hydrofining reactor and the hydrocracking reactor, thereby improving the product quality of high aromatic potential heavy naphtha, jet fuel, diesel and tail oil, reducing the light hydrocarbon yield, and effectively improving the economic benefits of the device. However, in this method, due to thermodynamic equilibrium limitations, the product quality improvement is limited. Summary of the invention

[0008] In view of the deficiencies in the prior art, the object of the present invention is to provide a wax oil hydrogenation treatment method, which can enable the raw oil to undergo a targeted hydrogenation reaction in a reactor by timely removing and separating the reaction products and setting the catalyst grade according to different components, thereby solving the problem of limited blending ratio due to the inability to better control the hydrogenation depth of catalytic diesel and wax oil raw materials, greatly increasing the aromatic content in heavy naphtha in the hydrocracking product, and improving the product quality of diesel and tail oil.

[0009] The present invention provides a wax oil hydrotreating method, comprising the following steps: (1) After the wax oil and the blended catalytic diesel are heated, they enter the middle of the hydrogenation pretreatment reactor together. A cavity section is set in the middle of the reactor. After flash evaporation, the light component enters the upper hydrogenation catalyst upward, and the heavy component enters the lower hydrogenation catalyst downward. Hydrogen enters from the bottom of the reactor and moves upward to the heavy component of the crude oil flowing down from the middle to perform film contact hydrogenation. (2) In step (1), the distillate oil after the heavy fraction is subjected to hydrogenation pretreatment enters the middle part of the hydrocracking reactor, the light fraction after the distillate oil is flashed upwardly enters the upper cracking catalyst bed layer, the number of the beds is 1 to 3 layers; the heavy fraction enters the lower cracking catalyst bed layer downwardly, the number of the lower cracking catalyst beds is 2 to 8 layers; wherein the average pore size of the lower cracking catalyst is 1 to 5 nm larger than the average pore size of the upper cracking catalyst; (3) The light components flowing upward in step (1) are pre-treated by hydrogenation and then flow out from the top of the reactor. After condensation, a high-aromatic naphtha product is obtained. The non-condensable gas can be used as circulating hydrogen after selective desulfurization. (4) The heavy fraction described in step (2) is respectively ring-opened and broken into naphtha fraction, middle distillate oil, and hydrogenated saturated unconverted oil under the action of a cracking catalyst and hydrogen; (5) The naphtha fraction and the intermediate oil fraction obtained in step (4) flow upwardly out of the top of the reactor together with hydrogen for condensation to obtain a high-aromatic naphtha fraction. The non-condensable gas can be used as circulating hydrogen after selective desulfurization. The intermediate distillate oil obtained in step (4) is cracked and ring-opened again when it flows upwardly through the cracking catalyst bed again to obtain a naphtha fraction.

[0010] Furthermore, the method further comprises step (6): the unconverted oil obtained in step (4) is stripped in a gas stripping tower, and the heavy fraction is sent out as an ethylene cracking product or returned to the feed inlet of a hydrocracking reactor for further cracking; the light fraction is condensed to obtain a naphtha fraction oil product, and the non-condensable gas is selectively desulfurized and used as circulating hydrogen.

[0011] The invention uses wax oil and catalytic diesel as raw materials, and mainly produces naphtha containing high aromatics and saturated straight-chain alkane chemical materials, so as to solve the problems of non-uniform cracking and hydrogenation depth and limited blending ratio caused by excessive difference in fraction composition in the existing wax oil blending and catalytic diesel refining process, further increase the aromatic content in naphtha, and fully digest the rich diesel intermediate fractions in the refinery.

[0012] The wax oil in step (1) of the present invention is straight-run wax oil, vacuum wax oil, coking wax oil, etc., and the catalytic diesel can be full-fraction catalytic diesel, catalytic light diesel or catalytic heavy diesel.

[0013] In the present invention, both the hydrogenation pretreatment reactor and the hydrocracking reactor are fed from the middle, and a light-heavy separation chamber is arranged in the middle of the hydrogenation pretreatment reactor and the hydrocracking reactor, and the volume V0 of the separation chamber is 0.5% to 30% of the effective volume V of the hydrogenation pretreatment reactor and the hydrocracking reactor, respectively, preferably 1% to 5%.

[0014] Furthermore, the separation chamber can be a cavity or a packing or a tower plate can be placed. The packing or tower plate are all conventional forms in the art, such as the packing can be selected from one or more random packings such as ball rings, Raschig rings, rectangular saddle rings, saddles, open-hole ring types, semi-rings, step rings, double arcs, Haier rings, conjugate rings, flat rings, and wreaths. The packing can also be selected from metal or ceramic corrugated packing. The tower plate is one or more of the tower plates with downcomers such as bubble plates, sieve plates, float valve plates, mesh plates, tongue-shaped plates, guide sieve plates, and multi-downcomer tower plates, and can also be tower plates without downcomers such as flow-through sieve plates and flow-through corrugated plates. High-efficiency tower plates such as float valve plates and sieve plates are preferred.

[0015] One to three layers of catalyst are placed on the top of the hydrogenation pretreatment reactor. When more than two layers of catalyst are placed, the hydrorefining catalyst and the hydrocracking catalyst are loaded from bottom to top. The volume and number of layers of the two catalysts can be determined according to the process conditions.

[0016] Furthermore, in the above technical solution, at least two layers are provided in the lower bed of the hydroprocessing pretreatment reactor described in step (1). Hydroprocessing catalyst and medium oil type cracking catalyst are sequentially loaded from top to bottom. The activity of the hydroprocessing catalyst increases sequentially from top to bottom of the bed, and the activity of the catalyst at the lowest end is lower than that of the upper bed. Similarly, the activity of the medium oil type cracking catalyst increases sequentially from top to bottom of the bed, and the activity of the catalyst at the lowest end is lower than that of the upper bed.

[0017] Furthermore, in the above technical solution, in step (1), the operating conditions of the upper bed of the hydrogenation pretreatment reactor are as follows: the average reaction temperature is 200-380°C, the reaction pressure is 8-20 MPa, the hydrogen-oil volume ratio is 200:1-2500:1, and the volume space velocity is 0.2-10.0 h -1 , the top reflux ratio is 1.2 to 4.5. The preferred operating conditions are: average reaction temperature is 260°C to 360°C, volume space velocity is 0.5 to 5.0 h -1 , the reaction pressure is 6 to 18 MPa, the hydrogen-to-oil volume ratio is 400: 1 to 1200: 1, and the top reflux ratio is 1.5 to 3.0.

[0018] In the above technical scheme, in step (1), the operating conditions of the lower bed of the hydrogenation pretreatment reactor are as follows: the average reaction temperature is 250-520°C, the reaction pressure is 8-20 MPa, the hydrogen-to-oil volume ratio is 200:1-2500:1, and the volume space velocity is 0.2-10.0 h -1 The preferred operating conditions are: average reaction temperature of 300°C to 400°C, volume space velocity of 0.3 to 3.0 h -1 , the reaction pressure is 6-18MPa, and the hydrogen-oil volume ratio is 500:1-1500:1.

[0019] Furthermore, in the above technical solution, in step (1), the temperature of the flash evaporation of the feedstock oil in the middle of the hydrogenation pretreatment reactor is 300°C to 450°C, preferably 320°C to 400°C.

[0020] In the hydrogenation raw materials, the sulfides exist mainly in the form of organic compounds such as hydrogen sulfide, mercaptan, sulfide, disulfide, and thiophene. Hydrogen sulfide and mercaptan are mainly present in distillate oils below 300°C, sulfide and disulfide are mainly present in middle distillate oils (200°C to 350°C), and sulfide, disulfide, and thiophene are in distillate oils above 350°C. In the present invention, the sulfides in wax oil are mainly sulfides and thiophenes, and the sulfides of straight-run wax oil include chain sulfides and cyclic sulfides. The sulfides in VGO are mainly cyclic sulfides. Thiophene sulfides are alkylthiophene, alkylbenzothiophene, alkyldibenzothiophene, etc. And with the increase of boiling point, the thiophene sulfur content first increases and then decreases, and the sulfides are concentrated in the 380°C to 530°C fraction. Monocyclic aromatic hydrocarbons (benzene, toluene, xylene, ethylbenzene, alkylbenzenes, phenylcycloalkanes) are mainly concentrated in the fractions with an initial boiling point of ~290℃, and the fractions with a temperature of >290℃ are mainly dicyclic and tricyclic aromatic hydrocarbons. The distribution of nitrogen is similar to that of sulfides and aromatic hydrocarbons. In the fractions with a temperature of less than 320℃, the nitrogen content only accounts for about 25% of the total fraction, of which the main nitrogen compounds are aniline and indole. The nitrogen content in the fractions with a temperature of more than 320℃ is mainly carbazole, accounting for about 75% of the total fraction. The order of hydrodenitrogenation of the three types of nitrogen compounds from easy to difficult is aniline>indole>carbazole.

[0021] Based on the above principle, the reactor of the present invention can have a temperature distribution ranging from 200°C to 520°C at different reactor positions through the setting of internal components and the adjustment of process parameters. Therefore, different special catalysts can be placed at different reactor positions according to the temperature distribution to carry out targeted hydrogenation treatment of different forms of sulfur, nitrogen and aromatics in respective fractions, thereby improving the conversion efficiency and service life of the catalyst.

[0022] In the present invention, the raw oil is flashed under the conditions described in the present invention, and under the action of hydrogen stripping, the distillate oil with a temperature less than 300°C is sent to the upper part of the hydrogenation pretreatment reactor. This part of the distillate oil mainly contains mercaptan sulfides, monocyclic aromatic hydrocarbons and aniline indole nitrogen compounds, which are relatively easy to remove. Therefore, the reaction can be carried out under relatively low temperature conditions, and the volume space velocity is also higher than that of conventional hydrofining reactors. After desulfurization and denitrification, the light component still has a small amount of monocyclic aromatic hydrocarbons and bicyclic aromatic hydrocarbons with long side chains and branches. After passing through the cracking catalyst described in the present invention, the side chains or branches can be opened, and the bicyclic rings can be opened to obtain lighter monocyclic aromatic hydrocarbons, further improving the light aromatic hydrocarbon content of the light oil product, and the moderate cracking depth is achieved by controlling the volume space velocity and the reflux ratio of the tower top.

[0023] In the present invention, the heavy fraction with a temperature greater than 300° C. enters the lower hydroprocessing catalyst bed layer. The lower bed layer controls a lower volume space velocity and a higher temperature, which can improve the removal rate of thiophene and carbazole impurities. The middle oil type cracking catalyst is placed at the bottom to further improve the yield of the middle distillate oil.

[0024] Furthermore, in the above technical solution, the hydrocracking reactor in step (2) is fed in the middle. The upper catalyst bed of the hydrocracking reactor is provided with 1 to 3 layers, and the lower catalyst bed is provided with 2 to 8 layers, and each layer can be provided with catalysts of different activities and volumes.

[0025] Furthermore, the average activity and average pore size of the cracking catalyst in each bed layer at the upper part of the cracking reactor are smaller than the average activity and average pore size of the catalyst in each bed layer at the lower part of the cracking reactor.

[0026] Furthermore, the catalytic activity of each bed layer in the lower part of the cracking reactor is preferably X1<X2<……<Xn (X represents activity, 1……n represents catalyst bed number), where catalyst activity is defined as the conversion rate of the catalyst to the raw material; the pore size of the catalyst in each bed layer in the lower part is preferably d1<d2<……<dn.

[0027] Furthermore, in the above technical scheme, the hydroprocessing catalyst described in steps (1) and (2) can be any suitable hydroprocessing catalyst available, preferably a hydroprocessing catalyst with high desulfurization and denitrification activity and low aromatic saturation. The hydrogenation active component can be selected from one, two or more metals in Group VIB and Group VIII, such as W, Mo, Co, and Ni, and the total content of the metal oxide is 5% to 75% by weight of the metal oxide. The carrier of the hydrogenation catalyst can be selected from alumina, amorphous silica-alumina, silicon oxide, titanium oxide, etc. At the same time, a part of the additives, such as P, Ti, Zr, Si, B, etc., can also be added to the catalyst. The hydroprocessing catalyst of the present invention can be a commercial catalyst, such as 3936, 3996, CH-20, FF-14, FF-16, FF-18, FF-24, FF-26, FF-34, FF-46, FF-56, FF-66, FH-UDS series, FZC-41, FZC-42 and other hydrogenation catalysts developed by Fushun Research Institute of Petrochemical Industry (FRIPP), HC-P, HC-T, HC-K developed by UOP, UF-210 / 220 and other hydrogenation catalysts, HR-406, HR-416, HR-448 and other hydrogenation catalysts from IFP, ICR154, ICR174, ICR178, ICR179 and other hydrogenation catalysts from CLG, TK-525, TK-555, TK-557 and other hydrogenation catalysts from Topsoe, KF-752, KF-756, KF-757, KF-840, KF-848, KF-901, KF-907 and other hydrogenation catalysts from AKZO, but not limited to the above catalysts.

[0028] Furthermore, in the above technical scheme, the medium oil type cracking catalyst carrier loaded in the lower part of the hydrogenation pretreatment reactor described in step (1) of the present invention includes β-type molecular sieve, Y-type molecular sieve, amorphous silica-alumina and binder. Based on the weight of the carrier, the content of β-type molecular sieve is 2-20wt%, the content of Y-type molecular sieve is 10-30wt%, the content of amorphous silica-alumina is 20-60wt%, and the content of binder is 15-40wt%; the hydrogenation active metal is a metal of Group VIB and / or Group VIII, the Group VIB metal is preferably molybdenum and / or tungsten, and the Group VIII metal is preferably cobalt and / or nickel. Based on the weight of the catalyst, the content of Group VIB metal (calculated as oxide) is 10wt%-30wt% and the content of Group VIII metal (calculated as oxide) is 4wt%-15wt%, and the content of the carrier is 60.0%-86.0%. The specific surface area of ​​the catalyst is 300-600 m 2 / g, and the pore volume is 0.3~0.7mL / g.

[0029] Furthermore, in the above technical scheme, the other hydrocracking catalysts used in the steps (1) and (2) can be any suitable existing hydrocracking catalysts, and preferably the active components of the catalyst are one or more of the metals of Group VIB and / or Group VIII. The metals of Group VIB are generally Mo and / or W, and the metals of Group VIII are generally Co and / or Ni. The carrier components of the catalyst include one or more of alumina, silicon-containing alumina and molecular sieves, preferably molecular sieves, and the molecular sieves can be Y-type molecular sieves. In a preferred embodiment, based on the weight of the catalyst, it contains 10wt% to 35wt% of the Group VIB metal in terms of oxide, 3wt% to 15wt% of the Group VIII metal in terms of oxide, 5wt% to 40wt% of the molecular sieve, 15wt% to 72wt% of amorphous silica-alumina, and 10wt% to 67wt% of alumina; the specific surface area of ​​the catalyst is 100 to 650m 2 / g, pore volume is 0.15-0.50mL / g. The hydrocracking catalyst can be a commercial catalyst, and the main commercial catalysts include 3963, FC-18, FC-32, FC-60, FC-70, FC-76, FC-80 catalysts developed by Fushun Petrochemical Research Institute, HDK786, HDK776, HDK766 catalysts of Axens, HC215, HC-120, DHC-8, HC-150, HC-140, HC43LT, HC190, HC-185, HC-175, HC-24 and other catalysts of UOP.

[0030] Studies have shown that thin film contact between gas and liquid on the solid surface reduces the resistance of gas mass transfer to the solid catalyst surface, and the continuous removal of reaction products from the system allows the equilibrium reaction to continue to proceed to the right, greatly improving the reaction efficiency and speed. The use of gas-liquid raw materials in the bed in reverse contact, using the different boiling points of the components in the product, so that the light distillate oil in the product is continuously separated from the reaction system, not only reduces the probability of continued deep cracking of the light distillate oil, but also facilitates the forward progress of the main reaction, reducing the difficulty of controlling operating conditions and controlling the reaction process. At the same time, the reverse contact reactor can continuously remove the reaction heat from the catalyst system, and can also prevent adverse factors such as sudden rise in bed temperature and aggregation and coking of by-products.

[0031] Furthermore, in the above technical solution, the shape of the catalyst added in steps (1) and (2) can be a mixture of one or more of conventional hydrocracking catalysts, porous catalysts, and shaped catalysts, preferably a porous catalyst, a shaped catalyst, and a mixed catalyst with conventional catalysts. The pore size of the porous catalyst is 1 to 50 mm, preferably 4 to 20 mm; the average particle diameter of the shaped catalyst is 2 to 50 mm, preferably 4 to 30 mm; the diameter or side length of the honeycomb catalyst pores is 1 to 50 mm, preferably 3 to 15 mm; the recommended porosity of the catalyst bed is 15% to 85%, preferably 20% to 75%.

[0032] Furthermore, in the above technical scheme, the shape of the special-shaped catalyst is a honeycomb, ball ring, Raschig ring, rectangular saddle ring, saddle shape, open-hole ring type, half ring, step ring, double arc, Haier ring, conjugate ring, flat ring, wreath, hollow sphere, gear shape, clover shape, etc.

[0033] Furthermore, in the above technical solution, the extraction temperature of the top distillate oil from the hydropretreatment reactor and the hydrocracking reactor in step (3) and step (4) is 140 to 240°C, preferably 160 to 210°C.

[0034] Furthermore, in the above technical solution, 1-2 side lines can be arranged at the top of the middle cavity of the hydrocracking reactor in step (2). The extraction temperature of the first side line is 180-250°C, preferably 180-230°C, and the component extracted from the side line is a jet fuel fraction; the temperature of the second side line is 160-380°C, preferably 180-360°C, and the component extracted is a diesel component.

[0035] Furthermore, from the feedstock oil inlet to the unconverted oil outlet, catalysts with different functions can be placed in each sub-bed. Preferably, for the catalyst with hydrocracking function, the average pore size of the catalyst surface increases gradually.

[0036] Furthermore, in the above technical solution, the raw materials of low-quality diesel and hydrogen in step (1) are heated in a heating furnace before entering the hydroprocessing reaction zone, preferably heated to 200-350° C. The separation in step (1) is preferably carried out in a gas-liquid separation tank, and after separation, the gas (containing hydrogen and non-condensable gas) is pressurized from the top of the separation tank by a hydrogen compressor and enters the hydroprocessing reaction zone as circulating hydrogen.

[0037] Furthermore, in the above technical solution, the stripping tower described in step (4) is of conventional design in the art, and can be provided with a reboiler or steam stripping. The bottom temperature of the tower is generally 200 to 350°C, preferably 250 to 300°C. The condensation temperature at the top of the stripping tower is generally 140 to 240°C, preferably 160 to 210°C.

[0038] Compared with the prior art, the present invention has the following beneficial effects: 1. The present invention adopts a feeding method in the middle of the reactor. Under the action of the separation chamber, the light and heavy components in the raw oil enter the upper and lower parts of the reactor respectively, so that different raw oils can be screened into distillate oils in the reactor. The required hydrotreating or hydrocracking catalysts are loaded in the upper and lower parts of the reaction zone in a targeted manner. By changing the operating conditions, the hydrogenation depth of the distillate oil can be controlled in one reactor, and the naphtha yield is finally improved. At the same time, the polycyclic aromatic hydrocarbons in the raw material are moderately opened on the catalyst beds with different functions, thereby increasing the content of light aromatics in the light oil.

[0039] 2. The gas and liquid raw materials of the present invention realize a thin film reverse contact mass transfer reaction on the surface of the solid catalyst, which reduces the thickness of the traditional solid-liquid immersion mass transfer membrane, strengthens the mass transfer, and improves the reaction speed.

[0040] 3. In the present invention, the light components and reaction heat generated by the reaction can be removed from the reaction bed at any time, which is particularly suitable for improving the hydrogenation efficiency of hydrogenation exothermic and heat balance reactions. The continuous removal of light components greatly improves the yield of the target product.

[0041] 4. In the present invention, reaction and separation are concentrated in one reactor, and catalysts with different functions are also placed in the same reactor. Compared with the traditional multi-reactor and multi-fractionation tower process, it can save high-separator, low-separator, and supporting distillation towers and other equipment. The process is short and easy to operate, saving space and energy consumption.

[0042] 5. Compared with the existing hydrocracking technology, the catalytic diesel oil of the present invention can be blended with wax oil in any proportion, which improves the flexibility of the device. Under the same raw materials and process conditions, the yield of monocyclic aromatic hydrocarbons of the present invention is 2 to 8 percentage points higher than that of the conventional fixed bed two-stage process. The naphtha fraction obtained by the present invention is rich in aromatic hydrocarbons and is a high-quality raw material for the reforming device. The unconverted oil obtained is a high-quality ethylene cracking material. BRIEF DESCRIPTION OF THE DRAWINGS

[0043] Figure 1 This is a schematic flow diagram of a wax oil hydrogenation treatment method of the present invention; Among them: 1-raw oil inlet; 2-new hydrogen; 21, 22-circulating hydrogen; 3-hydrogenation pretreatment reactor; 4-pretreatment tower top condenser; 5-pretreatment tower top liquid separator; 6-pretreatment generated oil; 7-hydrogenation cracking raw oil; 8-hydrogenation cracking reactor; 9-cracking tower top condenser; 10-cracking tower top liquid separator; 11-cracking tower bottom reboiler; 12-unconverted oil; 13-gas stripping tower; 14-gas stripping tower top condenser; 15-gas stripping tower top partial liquid tank; 16-gas stripping tower bottom reboiler; 17-circulating oil; 18-ethylene cracking material; 19-intermediate distillate oil; 20-desulfurization and deamination equipment; 21-circulating hydrogen compressor; 22, 23-gas-liquid separation chamber; 24-naphtha fraction. DETAILED DESCRIPTION

[0044] The specific implementation modes of the present invention are described in detail below in conjunction with the accompanying drawings, but it should be understood that the protection scope of the present invention is not limited by the specific implementation modes.

[0045] Unless explicitly stated otherwise, throughout the specification and claims, the term “comprise” or variations such as “include” or “comprising”, etc., will be understood to include the stated elements or components but not to exclude other elements or components.

[0046] In this document, for the convenience of description, spatial relative terms such as "below", "below", "down", "above", "above", "upper", etc. may be used to describe the relationship between one element or feature and another element or feature in the accompanying drawings. It should be understood that the spatial relative terms are intended to include different directions of objects in use or operation in addition to the directions drawn in the drawings. For example, if the object in the figure is turned over, the elements described as being "below" or "below" other elements or features will be oriented "above" the elements or features. Therefore, the exemplary term "below" can include both below and above. Objects may also have other orientations (rotated 90 degrees or other orientations) and the spatial relative terms used in this document should be interpreted accordingly.

[0047] In this document, the terms "first", "second", etc. are used to distinguish two different elements or parts, and are not used to limit a specific position or relative relationship. In other words, in some embodiments, the terms "first", "second", etc. can also be interchangeable.

[0048] Figure 1 The figure shows a schematic diagram of the wax oil hydrogenation process of the present invention. Figure 1The process flow of the present invention includes: the wax oil mixed with catalytic diesel is heated to 200-380°C and then enters the separation chamber 23 in the middle of the hydrogenation pretreatment reactor, and is separated into light components and heavy components in the separation chamber 23. The light components are desulfurized, denitrogenated and hydrogenated upwards by the hydrotreatment catalyst, and then cooled by the pretreatment tower top condenser 4 and enter the liquid separator 5 for gas-liquid separation. The non-condensable gas rich in hydrogen passes through the desulfurization and deamination facility 20, and then returns to the bottom of the hydrogenation pretreatment tower by the circulating hydrogen compressor 21 as the device hydrogen 2-1 to enter the tower and enter the catalyst bed upward. The heavy components enter the hydrotreatment catalyst bed and the hydrocracking catalyst bed in turn downwards, and react with the hydrogen flowing upward from the bottom of the tower in the bed on the catalyst surface. The hydrogen sulfide, ammonia and light hydrocarbons generated by the heavy components in the hydrogenation process of the catalyst bed immediately enter the top of the reactor with the hydrogen, and no longer continue to undergo hydrogenation reaction with the heavy components. The generated oil 6 after hydrogenation pretreatment is mixed with the circulating oil 17 as the raw oil 8 and enters the middle separation chamber 23 of the hydrocracking reactor 8. Similarly, the light components separated in the separation chamber 23 enter the cracking catalyst bed upward, and the heavier components are selectively opened and saturated. The lighter components pass through the bed to the top condenser 9 for cooling. After gas-liquid separation in the liquid separator 10, the light components enter the circulating hydrogen system. The condensed liquid hydrocarbons are sent out of the separation chamber 23 as high aromatic products 24. The separated heavy components enter the catalyst bed downward and react with the hydrogen 2-2 entering from the top of the reactor in reverse film contact mass transfer to undergo hydrogenation and cracking reactions. The light components produced by the reaction are separated from the catalyst bed and finally condensed upward to form high aromatic products 24. The unconverted oil returns to the reactor inlet from the bottom of the tower as circulating oil 17, and can be used as a chemical material 18 to send out the device.

[0049] The following examples further illustrate a wax oil hydrogenation method and the effects of the invention. In the examples, detailed implementation methods and specific operation processes are given based on the technical solution of the present invention, but the protection scope of the present invention is not limited to the following specific examples. For example, various types of existing protective agents are placed in the hydrogenation pretreatment catalyst bed, which are all protected by the present invention.

[0050] The experimental methods in the following examples, unless otherwise specified, are conventional methods in the art. Example 1

[0051] Using the present invention Figure 1 In the flow chart shown, the feedstock is a mixture of vacuum wax oil and catalytic diesel, with the volume of catalytic diesel accounting for 27% (1# feedstock). The feedstock and hydrogen are heated to 400°C in a heating furnace and then enter the separation chamber of the hydrogenation pretreatment reactor together. The separation chamber accounts for 1% of the total effective volume of the reactor. Step ring packing is placed in the separation chamber, with 2 layers placed on the top. Catalyst A, catalyst B, V are loaded in sequence from bottom to top. A :V B=2:1. The lower part is filled with 3 layers of catalysts, from top to bottom, catalyst A, catalyst A, catalyst B, V A :V B =3:1. Hydrogen enters from the bottom of the reactor, passes through the lower catalyst bed and the upper catalyst bed in turn, and finally reaches the top of the reactor with light hydrocarbons, hydrogen sulfide, ammonia and other light components produced by hydrogenation, and is condensed and separated into light hydrocarbon products and hydrogen-rich non-condensable gas through the condenser. The heavy components after hydrogenation refining flow out from the bottom of the tower and enter the separation chamber in the middle of the hydrocracking reactor. The separation chamber accounts for 1.5% of the total effective volume of the reactor. The separated light components pass upward through the two upper cracking catalyst beds. The catalysts are loaded with catalyst C, catalyst D, and V in turn. C :V D =1.5:1. The separated heavy components go down to the catalyst beds of catalyst C, catalyst C, and catalyst D in turn. The light hydrocarbons and hydrogen produced after the reaction leave the bed together with the light components separated from the upper part and the light components produced by the hydrogenation of the bed, and are condensed and separated into light hydrocarbon products and non-condensable gases through the condenser at the top of the reaction zone. The unconverted oil at the bottom of the hydrocracking reactor enters the stripping tower to remove the light components in the unconverted oil and then send it out from the bottom of the stripping tower as a chemical cracking material. The light components at the top of the stripping tower are condensed and separated into high-aromatic light oil products and non-condensable gases. The non-condensable gases at the top of the hydrogenation pretreatment reaction zone, the hydrocracking reactor, and the stripping tower are mixed together and then recycled as circulating hydrogen after desulfurization and pressure increase. The condensed light oil products are sent out of the device together as products. Example 2

[0052] The difference between this embodiment and embodiment 1 is that the raw material oil used is 2# raw material oil (catalytic diesel accounts for 50% by volume), and two beds are set on the upper part of the hydrogenation pretreatment reactor, and catalyst A, catalyst A, and catalyst C are placed in sequence from the separation chamber upward. The other conditions are the same as those in embodiment 1. Example 3

[0053] The difference between this embodiment and embodiment 1 is that the hydrogenation pretreatment catalyst is A. The other conditions are the same as those in embodiment 1. Example 4

[0054] Same as Example 1, except that the catalyst in the hydrocracking reactor is D. The other conditions are the same as Example 1. Example 5

[0055] The same as Example 1, except that the catalyst used in the upper and lower catalysts of the hydrocracking reactor is Catalyst E, and the bed voidage is 30%.

[0056] The properties of the catalysts in the above-mentioned embodiments and comparative examples are shown in Table 1, the properties of the feedstock oil are shown in Table 2, the process conditions are shown in Table 3, and the product distribution of the embodiments and comparative examples is shown in Table 4.

[0057] Comparative Example 1 A conventional two-stage hydrogenation process, i.e., a refining + cracking process, is adopted. Both the refining and cracking reactors adopt a reaction process in which the raw material and hydrogen flow in parallel from top to bottom. The cracked distillate oil is then separated by a separation tower and sent out as a product. In the comparative example, the hydroprocessing reactor is filled with catalyst A and catalyst B, with a volume ratio of 2.5:1. The cracking reactor is filled with catalyst D and catalyst C from top to bottom in sequence. It is equipped with hot high fraction, cold high fraction, cold high fraction, cold low fraction, gas stripping tower, fractionation tower, and fractionation tower front heating furnace equipment. The hydrogenated product is fractionated by the fractionation tower to separate the light oil, and the tail oil at the bottom of the fractionation tower is sent out as a chemical cracking raw material. The remaining process conditions are the same as those in Example 1.

[0058] Table 1 Catalyst properties Catalyst A Catalyst B Catalyst C Catalyst D Catalyst E Catalyst carrier shape Raschig Ring Raschig Ring Raschig Ring Raschig Ring Cylindrical bar Catalyst inner ring diameter, mm 2.5 2.5 3.0 3.0 — Catalyst size, mm 4.5mm×4.5mm 4.5mm×4.5mm 5mm×5mm 5mm×5mm φ1.8mm / 5~8mm Bed void ratio after loading, % 70 70 73 73 35 Metal content <![CDATA[MoO3,wt%]]> 23.9 <![CDATA[WO 3， wt%]]> 21.9 15.3 20.5 15.1 NiO, wt% 4.1 5.8 6.5 7.3 6.4 <![CDATA[Specific surface area, m 2 / g]]> 310 410 690 681 420 Pore ​​volume, ml / g 0.32 0.51 0.48 0.42 0.41 Average pore size, nm 9 9 6 8 Table 2 Properties of crude oil Crude oil Vacuum wax oil Catalytic diesel Wax oil + 27% fuel oil Wax oil + 50% accelerator serial number 1# 2# <![CDATA[Density (20 °C), g / cm 3 > 0.918 0.949 0.924 0.935 Distillation range, ℃ 292～515 174～357 203～520 196～523 Aromatics, wt% 37.9 78.2 48.5 58.2 (S)% 2.6 1.1 2.18 1.78 <![CDATA[ w (N / μg·g -1 ]]> 850 6100 2315 3200 Table 3 Process conditions of the embodiments and comparative examples Hydrogenation pretreatment reactor Reactor inlet temperature, °C 400 Average reaction temperature of upper bed, ℃ 345 Average reaction temperature of lower bed, ℃ 368 Reaction pressure, MPaG 14.0 Inlet volume hydrogen to oil ratio 800：1 <![CDATA[Space velocity, h -1 > 1.0 Hydrocracking reactor Average reaction temperature of upper bed, ℃ 365 Average reaction temperature of lower bed, ℃ 395 Reaction pressure, MPaG 14.0 Inlet volume hydrogen to oil ratio 1200:1 <![CDATA[Space velocity, h -1 > 1.5 Table 4 Product distribution of embodiments and comparative examples

[0059] From the results in Table 4, it can be seen that the method of the present invention, since the light and heavy fractions enter different catalyst beds for hydrogenation treatment in the reactor, the raw materials are selected from the process, the division of labor of the catalysts is achieved, and the light oil yield is effectively improved; at the same time, the aromatic content in the light oil is greatly improved. Under the same raw materials, the light oil yield of <210°C is increased by 9.4 percentage points, and the aromatic yield is increased by 8 percentage points. Under the same cracking catalyst, the light oil yield is increased by 3.1 percentage points, and the aromatic yield is increased by 2.1 percentage points.

Claims

1. A wax oil hydrotreating method, characterized in that: The steps include: (1) After being heated, wax oil and blended catalytic diesel enter the middle of a hydrogenation pretreatment reactor together, and a cavity section is arranged in the middle of the reactor; after flash evaporation, the light component enters the upper hydrogenation catalyst upward, and the heavy component enters the lower hydrogenation catalyst downward; hydrogen enters from the bottom of the reactor and moves upward to perform film contact hydrogenation with the heavy component of the crude oil flowing down from the middle; (2) In step (1), the distillate oil after the heavy component is subjected to hydrogenation pretreatment enters the middle part of the hydrocracking reactor, the light fraction after the distillate oil is flashed enters the upper cracking catalyst bed upward, and the heavy fraction enters the lower cracking catalyst bed downward; wherein the average pore size of the lower cracking catalyst is 1 to 5 nm larger than the average pore size of the upper cracking catalyst; (3) The light components flowing upward in step (1) are pre-treated by hydrogenation and then flow out from the top of the reactor, and are condensed to obtain a high-aromatic naphtha product; (4) The heavy fraction described in step (2) is respectively ring-opened and broken into naphtha fraction, middle distillate oil, and hydrogenated saturated unconverted oil under the action of a cracking catalyst and hydrogen; (5) The naphtha fraction and the intermediate oil fraction obtained in step (4) flow upward out of the top of the reactor together with hydrogen for condensation to obtain a high-aromatic naphtha fraction. The intermediate distillate oil obtained in step (4) is cracked and ring-opened again when it flows upward through the cracking catalyst bed again to obtain a naphtha fraction.

2. The wax oil hydrotreating method according to claim 1, characterized in that: The method further comprises step (6): the hydrogenated saturated unconverted oil obtained in step (4) is stripped in a gas stripping tower, and the heavy fraction is sent out as an ethylene cracking product or returned to the feed inlet of the hydrocracking reactor for further cracking; the light fraction is condensed to obtain a naphtha fraction oil product.

3. The wax oil hydrotreating method according to claim 1, characterized in that: The wax oil is selected from at least one of straight-run wax oil, vacuum wax oil and coking wax oil.

4. The wax oil hydrotreating method according to claim 1, characterized in that: A light-heavy separation chamber is arranged in the middle of the hydrogenation pretreatment reactor and the hydrocracking reactor, and the volume V0 of the separation chamber is 0.5% to 30% of the effective volume V of the hydrogenation pretreatment reactor and the hydrocracking reactor respectively.

5. The wax oil hydrotreating method according to claim 4, characterized in that: The separation chamber is a hollow cavity or is provided with fillers or tower plates.

6. The wax oil hydroprocessing method according to claim 1, characterized in that: More than two layers of catalysts are placed on the upper part of the hydrogenation pretreatment reactor, and the hydrorefining catalyst and the hydrocracking catalyst are sequentially loaded from bottom to top.

7. The wax oil hydroprocessing method according to claim 1, characterized in that: At least two catalyst beds are arranged at the lower part of the hydrogenation pretreatment reactor.

8. The wax oil hydrotreating method according to claim 7, characterized in that: The hydrotreating catalyst and the medium oil type cracking catalyst are loaded in sequence from top to bottom; the activity of the hydrotreating catalyst increases in sequence from top to bottom of the bed, and the activity of the lowest end catalyst is lower than that of the upper bed; the activity of the medium oil type cracking catalyst increases in sequence from top to bottom of the bed, and the activity of the lowest end catalyst is lower than that of the upper bed.

9. The wax oil hydroprocessing method according to claim 1, characterized in that: The operating conditions of the upper bed of the hydrogenation pretreatment reactor are as follows: average reaction temperature is 200-380°C, reaction pressure is 8-20MPa, hydrogen-oil volume ratio is 200:1-2500:1, volume space velocity is 0.2-10.0h -1 , the top reflux ratio is 1.2~4.

5.

10. The wax oil hydroprocessing method according to claim 1, characterized in that: The operating conditions of the lower bed of the hydrogenation pretreatment reactor are as follows: average reaction temperature of 250-520°C, reaction pressure of 8-20 MPa, hydrogen-oil volume ratio of 200:1-2500:1, volume space velocity of 0.2-10.0 h -1 .

11. The wax oil hydroprocessing method according to claim 1, characterized in that: In step (1), the temperature of the flash evaporation of the crude oil in the middle of the hydrogenation pretreatment reactor is 300°C to 450°C.

12. The wax oil hydroprocessing method according to claim 1, characterized in that: The hydrocracking reactor has 1 to 3 catalyst beds on the top and 2 to 8 catalyst beds on the bottom, with catalysts of different activities and volumes placed on each layer; the average activity and average pore size of the cracking catalysts in each bed on the top of the hydrocracking reactor are smaller than those in each bed on the bottom of the hydrocracking reactor.

13. The wax oil hydroprocessing method according to claim 1, characterized in that: In step (2), 1-2 side lines are arranged at the top of the middle cavity of the hydrocracking reactor.

14. The wax oil hydrotreating method according to claim 13, characterized in that: The extraction temperature of the first side line is 180-250°C, and the component extracted from this side line is a jet fuel fraction; the temperature of the second side line is 160-380°C, and the component extracted is a diesel component.

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