Hydrogenation treatment method and device for producing high-aromatic naphtha and BTX from inferior diesel

Abstract

The present invention discloses a method and device for producing high-aromatic naphtha and BTX from inferior diesel. The crude oil enters the hydrogenation pretreatment reactor and reacts with the upward hydrogen in the hydrogenation pretreatment bed; the hydrogenation pretreatment oil enters the hydrogenation reforming reactor, and the heavy component and hydrogen react in countercurrent contact in the multi-bed hydrogenation reforming reaction zone, and the light product obtained by the lighter part and cracking enters the distillation chamber upward, and the naphtha product is obtained in the side line. The method of the present invention allows the light component to quickly leave the system and no longer excessively participate in saturation and cracking. Since the product can quickly leave the reaction system, while increasing the positive reaction speed, it also eliminates the hidden danger of by-products clogging the catalyst, thereby improving the yield of the target product.

Description

Technical Field

[0001] The invention belongs to the field of hydrogenation reaction, and in particular relates to a method and a device for hydrogenation treatment of inferior diesel. Background Art

[0002] With the improvement of people's quality of life and the stricter global environmental protection requirements in the future, the proportion of new energy vehicles is getting higher and higher. Instead, the demand for gasoline and diesel is getting lower and lower, and the market competition in this part is becoming more and more brutal. In particular, the demand for diesel has gradually declined since 2016. At the same time, new large-scale oil refineries are constantly increasing, diesel production capacity is constantly increasing, diesel product surplus is constantly intensifying, and the competitive environment is deteriorating. Among diesel, catalytic diesel accounts for the largest proportion, but the quality is the worst, with an aromatic content of 60% to 85%. Generally, the excess catalytic diesel is reduced by modifying catalytic diesel to produce more gasoline or light aromatics (BTX).

[0003] The FD2G technology developed by Fushun Petrochemical Research Institute uses a two-stage reactor. The catalytic diesel first enters the hydrotreating reactor. After impurities are removed and moderate hydrogenation is performed, it is passed through a special strong cracking and weak hydrogenation light oil catalyst to control the depth of the hydrogenation reaction to avoid hydrogenation saturation of monocyclic aromatics, retain more monocyclic aromatics, and partially convert low cetane diesel into high-octane gasoline or aromatic raw materials. The research octane number (RON) of the gasoline product using this process is 90.7-94.6%, and the BTX mass fraction is 20.29%-31.57%.

[0004] The RLG technology of Sinopec Research Institute of Petroleum Processing (RIPP) uses the same two-stage process, with RN-411 as the hydrofining catalyst and RHC-100 with a molecular sieve channel structure and a gradient distribution of acidity as the hydrocracking catalyst. When this technology processes catalytic diesel with an aromatic mass fraction of 74%, the gasoline yield is 38.44% and the RON reaches 93.5.

[0005] The LTAG combined process developed by RIPP uses selective hydrogenation-catalytic cracking to produce high-octane gasoline or light aromatics, which can increase BTX production while reducing hydrogen consumption. Chiyoda Corporation of Japan has announced a non-hydrogenated aromatics production technology (FCA), which does not require hydrogenation saturation, but directly produces light aromatics such as BTX through fluidized catalytic cracking, with a BTX yield of up to 35%.

[0006] UOP's LCO-X process and Sinopec Shanghai Research Institute of Chemical Industry's PAC platform technology hydroprocessing + hydrocracking + aromatics maximization route. The LCO-X process combines the LCO-Unicracking process with the aromatics unit to convert the supersaturated alkyl cyclohexane and unconverted non-aromatic substances in the hydrocracking section into aromatics and LPG. The LCO-A technology developed by the Shanghai Institute processes 34.2% of the volume fraction of light aromatics in the LCO product.

[0007] Patent CN201810948782.X developed a catalyst using modified Y molecular sieve as a new active component. When used to process catalytic cracking diesel LCO, the BTX content in the gasoline product is 37.7%.

[0008] The above technologies all have the problem of over-hydrogenation or insufficient hydrogenation of the distillate oil in the hydrocracking unit, which makes the aromatics in LCO over-saturated to cycloalkanes, or only open the ring to dicyclic aromatics, causing the BTX yield in the product to reach a bottleneck value and cannot be further improved.

[0009] Patent CN201711470751.X first fractionates diesel raw materials into light fractions and heavy fractions. After aromatics extraction, the light fraction is obtained to obtain light aromatic fractions and light non-aromatic fractions; the heavy fraction and light non-aromatic fraction are successively subjected to hydrofining and light reforming reactors to finally obtain a gasoline component with a high aromatic content. This patent improves process flexibility and product quality by performing targeted hydrogenation on different components. However, the above process is relatively complicated. In addition, the hydro-reforming reactor and the hydro-refining reactor use traditional fixed-bed reactors, and there is a situation where the distillate oil is over-hydrogenated or cracked, which ultimately affects the light aromatics yield.

[0010] In summary, although existing technicians have developed different processes and catalysts in an effort to achieve a higher BTX yield, they are limited by the inherent defects of fixed bed reactors, namely, since the cracking products may undergo cracking reactions again during the cracking process, the reaction depth is uncontrollable. In this reaction system, the selectivity of distillates such as light and heavy naphtha is limited, and the reactor will be coked and blocked. In addition, the hydrogenation reactor also has the risk of "heating up" or flooding of the catalyst bed, especially for catalysts with higher activity. The reactor temperature rises at an alarming rate, usually reaching 500°C to 600°C, and some exceed 800°C. In severe cases, the catalyst may fail and be scrapped.

[0011] The height-to-diameter ratio of the fixed bed in traditional hydrogenation reactions (the ratio of the total height of the reactor bed to the diameter) is generally selected to be 2 to 10 to ensure sufficient contact between the gas-liquid material and the solid catalyst to achieve the required reaction depth and efficiency. Dong Fangliang et al. mentioned in "Determination of the Main Structural Parameters of the Hydrogenation Reactor" in "First Heavy Technology" 1998.1 (Total 75) that in order to avoid the small height-to-diameter ratio "causing uneven fluid distribution and poor catalyst contact efficiency", "the height-to-diameter ratio of the traditional fixed bed reactor is usually 4 to 9". Patent CN109679689 A also mentioned that the height-to-diameter ratio of the existing liquid phase hydrogenation reactor is generally 2.5 to 12. The design of the height-to-diameter ratio of the bed of the above-mentioned hydrogenation reactor has become a solidified cognition of technicians in this field. The application of a large number of industrial practices has also proved that the design is reasonable and has a relatively general adaptability. The extensive industrial success may also lead to the technicians' failure to conduct a more comprehensive and in-depth study of whether there are other better options for different types of reactions. For a long time, there have been no relevant research reports, or only research reports that prove that the existing height-to-diameter ratio is a suitable design. Summary of the invention

[0012] In view of the shortcomings of the prior art, the purpose of the present invention is to provide a method and device for hydrogenating inferior diesel to produce more BTX or naphtha, which can effectively control the reaction depth to achieve the purpose of increasing the yield of BTX products, while eliminating the risk of bed "flying temperature" or flooding and extending the service life of the catalyst.

[0013] The present invention provides a hydrogenation treatment method for producing more BTX from inferior diesel, comprising the following steps:

[0014] (1) After the low-quality diesel fuel enters the liquid distribution assembly of the hydrotreating reactor and is dispersed, the lighter part is carried upward by the hydrogen flowing upward from the bottom, and the heavier part enters the hydrotreating bed downward to generate hydrotreating oil;

[0015] (2) The light components described in step (1) are separated upward in the separation zone of the hydrogenation pretreatment reactor, and the naphtha product is extracted from the side line;

[0016] (3) The hydropretreatment product oil obtained in step (1) enters the liquid distributor of the hydroprocessing reactor, and the light components in the distillate oil dispersed by the distributor enter the separation section upward; the heavy components enter the multi-bed hydroprocessing reaction zone downward;

[0017] (4) Every time the heavy component in step (3) enters a catalyst bed, the light component produced enters the separation section in step (3) in a timely manner; and finally, the unreacted distillate oil is fully circulated and enters the inlet of the hydro-reforming reactor;

[0018] (5) After the light components described in step (3) are separated in the separation section, a naphtha fraction is extracted from the side line and mixed with the naphtha fraction obtained in step (2).

[0019] Furthermore, the method further comprises step (6): the naphtha fraction product described in the above step (5) enters an aromatic extraction unit to separate light aromatic BTX products and light non-aromatic fraction oil.

[0020] Furthermore, the light non-aromatic fraction oil obtained in the above step (6) enters a supplementary hydrotreating reactor, and the refined oil is used as ethylene cracking feed.

[0021] Furthermore, in the above technical solution, the hydrogenation pretreatment reaction zone in step (1) contains at least one catalyst bed, and the ratio of the equivalent diameter of the cross-sectional area of ​​the catalyst bed to the total height of the catalyst bed (hereinafter referred to as the equivalent diameter-to-height ratio) is 2:1 to 10:1, preferably 3:1 to 6:1. Unless otherwise specified herein, the equivalent diameter-to-height ratio refers to the ratio of the equivalent diameter to the total height of the catalyst bed in the reactor. When there are multiple catalyst beds, the total height of the catalyst bed refers to the sum of the multiple catalyst beds.

[0022] Through research, it is found that when the overall space velocity is similar, for the gas-liquid-solid three-phase hydrogenation reaction in which the amount of liquid phase decreases rapidly and the amount of gas phase increases rapidly during the reaction process, when hydrogen is contacted with the raw oil in a gas-liquid countercurrent manner, the diameter-to-height ratio of the catalyst bed in the reactor is significantly higher than that of the existing conventional technology, so that the generated gas phase quickly leaves the catalyst bed, the adverse effects of the generated gas phase accumulate to a small extent, and the liquid phase has a more sufficient chance to react on the catalyst, thereby overcoming the traditional understanding that a small height-to-diameter ratio will bring about adverse effects such as poor contact effect, and achieving the effect of significantly improving the yield of the target product (the naphtha fraction in the present invention), while solving the problems of easy flooding of the countercurrent reactor and limited hydrogen-to-oil ratio.

[0023] Furthermore, in the above technical solution, the liquid in step (1) is dispersed into small droplets after passing through the liquid distribution component. Under the action of hydrogen stripping, the lighter part is taken away upward, and the heavier part enters the hydrogenation pretreatment bed layer under the action of gravity. The lighter part generally refers to the naphtha fraction, and the heavier part generally refers to the diesel fraction. In the catalytic cracking light cycle oil (LCO), the total aromatic content is as high as 80%, and the rest are alkanes, cycloalkanes and olefins, among which naphthalene-based bicyclic aromatics account for about 70% of the total aromatic content, and monocyclic aromatics and tricyclic aromatics each account for about 15% of the total aromatic content. The sulfur content of LCO is 0.2~1.5wt%, and the nitrogen content is 100~750ppm. Therefore, light aromatics and lighter naphtha fractions are taken to the separation section together. Bicyclic aromatics and tricyclic aromatics enter the hydrogenation pretreatment bed layer together with the diesel medium and heavy fractions.

[0024] Furthermore, the operating conditions of the hydrogenation pretreatment process described in step (1) are as follows: the average reaction temperature is 200-450°C, the reaction pressure is 3-20 MPa, the tower top reflux ratio is 1.2-4.5, the hydrogen oil volume ratio is 100:1-3000:1. The feed oil flow rate and the reactor catalyst volume ratio are 0.1-10.0h -1 The preferred operating conditions are: average reaction temperature of 260°C to 420°C, reaction pressure of 4 to 18 MPa, top reflux ratio of 1.5 to 3.0, and hydrogen to oil volume ratio of 200:1 to 1500:1.

[0025] Furthermore, the separation section in step (2) is provided with a refining reaction bed layer, which can be used to remove sulfides in the light aromatic components.

[0026] Furthermore, the operating conditions of the separation section in step (2) are as follows: the reflux temperature at the top of the separation section is 50°C to 70°C, preferably 55°C to 65°C; the reflux ratio is 1.1 to 10.0, preferably 1.5 to 3.0; the side line extraction temperature is 150°C to 185°C, preferably 160°C to 170°C. The pressure of the separation section is 3 to 20 MPa, preferably 4 to 18 MPa.

[0027] Furthermore, in the above technical solution, the hydrogenation pretreatment reaction zone in step (1) is divided into a plurality of sub-reaction zones along the longitudinal direction, so that when the hydrogen rises to the bottom of the bed through the distributor, the diffusion area of ​​the hydrogen can cover the cross-sectional area of ​​the bed at the bottom of the sub-reaction zone, and each sub-reaction zone is provided with a separate hydrogen inlet.

[0028] Furthermore, two catalysts are placed in the hydrotreating reaction zone described in step (1). A hydrotreating catalyst is placed in the upper layer, and a hydrocracking catalyst is placed in the lower layer. The volume ratio of the hydrotreating catalyst to the hydrocracking catalyst is 1:1 to 10:1, preferably 3:1 to 5:1. By grading two different types of catalysts, the high-aromatic feedstock oil can be desulfurized, denitrogenated and deoxygenated, and after appropriate cracking, the polycyclic aromatic hydrocarbons in the feedstock oil are partially saturated, the long chains are broken, and the clean medium-molecular-sized distillate oil is converted into the feedstock for the subsequent hydrocracking reaction.

[0029] Furthermore, in the above technical solution, the hydrorefining catalyst can be any suitable hydrorefining catalyst, preferably a hydrorefining catalyst with high desulfurization and denitrification activity and less 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. The carrier of the hydrogenation catalyst can be selected from alumina, amorphous silicon aluminum, 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 hydrorefining 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, IFP's HR-406, HR-416, HR-448 and other hydrogenation catalysts, CLG's ICR154, ICR174, ICR178, ICR179 and other hydrogenation catalysts, Topsoe's TK-525, TK-555, TK-557 and other hydrogenation catalysts, AKZO's KF-752, KF-756, KF-757, KF-840, KF-848, KF-901, KF-907 and other hydrogenation catalysts. But not limited to the above catalysts.

[0030] Furthermore, in the above technical scheme, the hydrocracking catalyst in the hydropretreatment reaction zone and the hydroupgrading reaction zone can be any suitable hydrocracking catalyst, and the preferred catalyst active component is one or more of the Group VIB and / or Group VIII metals. The Group VIB metal is generally Mo and / or W, and the Group VIII metal is generally Co and / or Ni. The carrier component of the catalyst includes one or more of alumina, silicon-containing alumina and molecular sieves, preferably containing 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 silicon aluminum, and 10wt% to 67wt% of alumina; the specific surface area of ​​the catalyst is 100 to 650m 2 / g, and the 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, and FC-32 catalysts developed by Fushun Petrochemical Research Institute.

[0031] Furthermore, the operating conditions of the hydro-reforming reaction zone in step (3) are as follows: the average operating temperature is 230°C to 450°C, the operating pressure is 3MPa to 20MPa, the hydrogen-to-oil volume ratio is 200:1 to 4000:1, and the ratio of the feedstock oil flow rate to the reactor catalyst volume is 0.1h -1 ~10.0h -1 , the ratio of inlet circulating heavy oil to raw oil is 1:3-10:1, and the top reflux ratio is 1.1-10.0. The preferred operating conditions are: average operating temperature of 280°C-420°C, operating pressure of 4MPa-18MPa, hydrogen-to-oil volume ratio of 300:1-1800:1, top reflux ratio of 1.5-4.0, and bottom heavy oil to raw oil circulation ratio of 1:1-8:1.

[0032] Furthermore, the operating conditions of the separation section in the hydro-upgrading reactor in step (3) are as follows: the reflux temperature at the top of the separation section is 50°C to 70°C, preferably 55°C to 65°C; the reflux ratio is 1.1 to 10.0, preferably 1.5 to 3.0; the side line extraction temperature is 150°C to 185°C, preferably 160°C to 170°C. The pressure of the separation section is 3 to 20 MPa, preferably 4 to 18 MPa.

[0033] Furthermore, the number of reaction zone beds in the reaction section of the hydro-reforming reactor in step (3) is 1 to 30, preferably 2 to 10. The height of each bed is 50 to 500 mm, preferably 100 to 300 mm, and the diameter-to-height ratio is 2:1 to 20:1, preferably 3:1 to 10:1. The liquid feedstock oil is in thin-film contact with the catalyst. The catalyst height of different beds can be determined according to process requirements.

[0034] In step (4) of the present invention, when the heavy fraction passes through each bed layer from top to bottom, it is cracked into light BTX and other small molecular hydrocarbons, which quickly separate from the cross-layer and enter the separator upward. The heavy fraction moving downward gradually decreases, and finally the distillate oil that has not been ring-opened or chain-broken returns to the reactor inlet, ensuring the purpose of maximizing the production of light BTX.

[0035] In step (6) of the present invention, the aromatics extraction process can adopt any method suitable for extracting aromatics in the prior art, and can adopt a single extraction or a double extraction device. The extraction solvent can be one or more of sulfolane, furfural, N-methylpyrrolidone, dimethyl sulfoxide, etc., preferably sulfolane. The operating conditions of the extraction process are: temperature 60~200℃, preferably 75~180℃, operating pressure 0.1~1.0MPaG, preferably 0.15~0.55MPaG, solvent ratio 1.0~8.0, preferably 2.0~4.0.

[0036] The supplementary refining reactor of the present invention is a conventional fixed bed reactor, generally with a reaction pressure of 3.0MPa to 15.0MPa, a reaction temperature of 300°C to 430°C, and a liquid hourly volume space velocity of 0.2h -1 ~6.0h -1 , the volume ratio of hydrogen to oil is 100:1~2000:1.

[0037] The present invention also provides a hydroprocessing device for producing more BTX from inferior diesel, the device comprising:

[0038] (1) A hydrogenation pretreatment reactor, which is a special-shaped hydrogenation reactor, comprising a heavy oil tank, a reaction chamber and a separator which are sequentially connected from bottom to top, and the reaction chamber is sequentially provided with: a liquid distribution component, a catalyst bed and a hydrogen distribution chamber from top to bottom;

[0039] (2) A hydrogenation reforming reactor, which is a hydrogenation distillation reactor, comprising: a reaction chamber, in which a catalyst bed is arranged in an annular manner, and the liquid crude oil and the hydrogen feed are in countercurrent contact in the catalyst bed; a stripping chamber, which is located in the middle of the hydrogenation distillation reactor and is a gas phase channel that runs through the reactor from top to bottom, and the gas phase products after the hydrogenation reaction of each catalyst bed enter the gas phase channel under the action of stripping; a distillation chamber, which is arranged at the upper part of the hydrogenation distillation reactor, and the gas phase products are distilled and separated in the distillation chamber.

[0040] Furthermore, in the above hydrogenation pretreatment reactor, the reaction chamber is a horizontal storage tank, the axial direction of which is arranged in the transverse direction, and the two ends of the horizontal storage tank are provided with heads. Alternatively, the reaction chamber is a flat cylindrical tank, the axial direction of which is arranged in the longitudinal direction.

[0041] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the cross-section of the catalyst bed in the reaction chamber is rectangular or circular, and the bed height is 100 to 5000 mm, preferably 200 to 1000 mm.

[0042] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, a plurality of partitions are arranged in parallel in the vertical direction in the hydrogen distribution chamber and the catalyst bed in the reaction chamber, and the plurality of partitions divide the hydrogen distribution chamber into a plurality of air inlet units, and at least one hydrogen inlet is arranged at the bottom of each air inlet unit. A plurality of holes are distributed on each partition; the partition extends upward to the catalyst layer, the opening rate of the partition below the catalyst bed is less than 70%, and the opening rate of the partition in the catalyst layer is greater than 50%.

[0043] Furthermore, in the above hydrogenation pretreatment reactor, there are 1 to 3 hydrogen distributors corresponding to the bottom of the catalyst bed between every two adjacent partitions, and the distribution area of ​​hydrogen from all hydrogen distributors in the partition area when reaching the bottom of the regional bed should cover the entire regional bed bottom. Furthermore, the partition is annular or circular.

[0044] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the heavy oil bunker is arranged at the bottom center of the reaction chamber, and the heavy oil bunker is connected to the multiple air inlet units.

[0045] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the liquid distribution assembly in the reaction chamber includes a liquid distributor and a liquid distribution plate and a distribution cone arranged above the liquid distributor. It is used to disperse the liquid generated by the hydrotreatment reactor into small droplets. Under the action of hydrogen stripping, the lighter part is taken out upward and the heavier part enters the hydrocracking reaction zone. The lighter part is usually the naphtha fraction, and the heavier part is usually the diesel fraction.

[0046] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the shape of the liquid distribution plate of the special-shaped hydrogenation reactor is the same as the cross-section of the catalyst bed of the hydrogenation main reactor, and the area of ​​the liquid distribution plate is 10% to 100% of the cross-section of the catalyst bed, preferably 60% to 100%.

[0047] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, a plurality of first through holes are evenly arranged on the distribution plate, a first overflow ring is arranged around the first through hole, and an overflow portion is arranged on the outer edge of the distribution plate; the opening rate of the distribution plate is 5% to 90%, the diameter of the hole is 5 to 100 mm, and the height of the overflow ring is 1 to 30 mm.

[0048] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, a serration portion is provided on the inner side of the overflow ring, the serration portion is bent downward, and a guide groove is provided on the serration portion.

[0049] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the distribution cone is arranged at the upper center of the liquid distribution plate, the distribution cone is provided with a plurality of holes, and an overflow ring is arranged around the holes. The top angle of the distribution cone is greater than 90°, the opening rate of the distribution cone is 5% to 80%, the height of the overflow ring is 1 to 30 mm; the bottom area of ​​the distribution cone is 2% to 15% of the area of ​​the liquid distribution plate.

[0050] The hydrogen feed pipe at the bottom of the lower reaction section of the reaction chamber of the hydrogenation pretreatment reactor of the present invention is provided with multiple inlets, each hydrogen feed pipe corresponds to the catalyst bed area between the two partitions, so that hydrogen can pass upward through the top reaction zone after coming out of each distributor. There is at least one channel at the connection between the perforated baffle and the bottom of the main reactor.

[0051] Furthermore, in the above-mentioned hydrogenation pretreatment reactor, the separator includes a mixing section, a separation section and a stabilization section from bottom to top. The height of the mixing section is 25% to 40% of the total height of the separator, the height of the separation section is 50% to 65% of the total height of the separator, and the height of the stabilization section is 5% to 10% of the total height of the separator.

[0052] Furthermore, the hydro-reforming reactor described in the present invention is a "blade" type hydro-distillation reactor, and the reaction chamber in the hydro-distillation reactor can be provided with: a liquid phase feed distributor, which is arranged above the topmost catalyst bed layer and is an annular tube type or an annular belt type; a hydrogen feed distributor, which is arranged at the bottom of each catalyst bed layer and is an annular tube type or an annular belt type.

[0053] Furthermore, the hydrogen feed distributor in the reaction chamber of the hydro-modification reactor may be arranged in a relatively closed space.

[0054] Furthermore, a gas guide unit may be provided between the hydrogen feed distributor and the bottom of the catalyst bed in the reaction chamber of the hydrogenation reforming reactor, and the gas guide unit includes: a guide plate, which is provided at the bottom of the catalyst bed and is used to receive the liquid phase product after the hydrogenation reaction of the catalyst bed; a guide pipe, which is evenly arranged on the guide plate and guides the hydrogen feed upward to the catalyst bed, and a first cap is provided on the top of the guide pipe; an overflow weir, which is close to one side of the distillation chamber and is used to overflow the liquid phase product on the guide plate to the next catalyst bed.

[0055] Furthermore, a liquid redistribution plate having an annular shape and uniformly provided through holes may be provided below the overflow weir in the reaction chamber of the hydro-modification reactor.

[0056] Furthermore, the liquid redistribution plate in the reaction chamber of the hydro-modification reactor may include: an overflow ring, which is arranged around the top of the through hole and is used to uniformly guide the liquid phase product to the next catalyst bed; a baffle, which is arranged on the side away from the distillation chamber, and the upper edge of the baffle is higher than the overflow ring.

[0057] Furthermore, the overflow ring in the reaction chamber of the hydro-modification reactor comprises: a serration portion, which is bent downward from the outer edge of the through hole toward the center; and a drainage groove, which is arranged on the serration portion and opened along the center of the serration portion.

[0058] Furthermore, a conical gas-liquid guide plate may be provided in the stripping chamber of the hydro-reforming reactor, and the number and position of the gas-liquid guide plate are matched with the catalyst bed.

[0059] Furthermore, an air hole is opened on the conical surface of the conical gas-liquid guide plate in the distillation chamber of the hydrogenation reforming reactor, and a second cap is provided on the top of the air hole; the gas phase products after the hydrogenation reaction rise from the air hole and converge into the distillation chamber, and part of the heavy component liquid phase products drip onto the gas-liquid guide plate and reflux into the reaction chamber.

[0060] Furthermore, the distillation chamber of the hydro-reforming reactor may include a mixing section and a separation section, and a packing or a tray may be provided in the separation section. The height of the separation section is 80% to 95% of the total height of the distillation chamber, and the height of the mixing section is 5% to 20% of the total height of the separator.

[0061] Furthermore, the catalyst in the hydrogenation pretreatment reactor and the hydrogenation reforming reactor is preferably a porous special-shaped catalyst such as a clover, a Raschig ring or a honeycomb; the porosity of the catalyst bed is 15% to 85%, preferably 20% to 75%.

[0062] Furthermore, in the above-mentioned hydrogenation pretreatment reactor and hydrogenation reforming reactor, the ratio of the diameter of the upper separator or distillation chamber or the diameter (or maximum equivalent diameter) of the lower reaction chamber is 1:1.2 to 1:50, preferably 1:2 to 1:10. In the present invention, the diameter of the separator or distillation chamber at the upper part of the reactor is reduced, so that the light fraction load under high pressure is fully matched with the tower plate, and the tower plate separation efficiency is high.

[0063] Furthermore, in the hydro-reforming reactor, the number of catalyst beds is 1 to 20, preferably 3 to 15. The diameter-to-height ratio may be 3:1 to 6:1, and the liquid phase feedstock oil is in thin film contact with the catalyst.

[0064] The low-quality LCO feedstock hydrogenation pretreatment reactor of the present invention can directly separate the light aromatics in the feedstock by using a special-shaped hydrogenation reactor, thereby reducing the load of the reaction part of the special-shaped hydrogenation reactor. After desulfurization, denitrification and other impurities, other heavier distillates are subjected to hydrogenation saturation and cracking reactions of a cracking agent to generate a small amount of light aromatics and medium molecular distillates as feedstock for the hydrogenation reforming reactor.

[0065] The hydrogenation reforming reactor of the present invention adopts a "blade" type hydrogenation distillation reactor, and the light aromatic hydrocarbons in the distillate oil after hydrogenation pretreatment are separated in the separation section of the distillation chamber. After the heavy distillate oil enters the cracking reaction bed, the polycyclic aromatic hydrocarbons open the ring to produce monocyclic aromatic hydrocarbons, and leave the reaction bed in time and are separated through the distillation chamber. The unopened polycyclic aromatic hydrocarbons and heavy distillate oil then enter the next bed. Similarly, the light aromatic fraction generated by the reaction leaves the reaction system in time and is separated. The unreacted heavy components react in turn in the following multiple beds. In this way, the reaction products are not removed, which speeds up the cracking reaction rate. Since the distillation chamber only separates the light aromatic products, the unseparated fractions in the distillation chamber enter the catalyst bed again for reaction. Therefore, the present invention is more selective for light aromatics, and can improve the yield of light aromatics, and the load of the catalyst bed is also greatly reduced. According to the test results, the load of each catalyst bed can be reduced by 5% to 20%.

[0066] The catalyst bed of the hydrogenation pretreatment reactor and the hydrogenation reforming reactor in the present invention can also greatly increase the material flux passing through the bed at a higher diameter-to-height ratio, while reducing the accumulation of heat in the catalyst bed. At the same time, the generated by-products such as hydrogen sulfide and ammonia are quickly taken out of the reactor, and the blockage phenomenon caused by the by-products staying in the bed for too long, which is caused by the traditional reactor, will not occur.

[0067] In the present invention, the heat generated by the hydrogenation reaction can be fully taken away in time from the bed through the hydrogen stripping and light component gasification process, and can be provided for the heat used for the gasification and distillation of the light component, with obvious energy-saving effect.

[0068] The present invention can disperse the liquid after hydrogenation treatment into suitable small droplets through the design and arrangement of the liquid distribution component, and with the help of the stripping effect of hydrogen, the lighter part can be directly taken out and no longer enter the hydrocracking reaction zone to participate in the hydrocracking reaction. At the same time, the design of the liquid distribution component allows the heavy components entering the hydrocracking reaction zone to be distributed more evenly, solving the problem of poor contact of reactants on the catalyst bed under a large diameter-to-height ratio in traditional reactors. Using the hydrogenation reactor of the present invention, under the same process conditions and product index requirements, the temperature rise of the reaction bed is controlled to be smaller and the allowable feed load is larger. The improvement of the pore structure can improve the liquid flooding characteristics (no liquid flooding for a long time) and ensure good mass transfer performance. Finally, the product properties have good controllability.

[0069] Compared with the prior art, the present invention has the following beneficial effects:

[0070] 1. The octane number or light aromatic content yield of the light naphtha of the present invention is higher. Using inferior LCO as raw material, the aromatic yield is 3 to 10 percentage points higher than that of the conventional fixed bed two-stage process. When the aromatic content in the LCO raw material is 80%, the yield of light aromatics (BTX) in the naphtha product reaches 43%; using ordinary inferior diesel as raw material (aromatic content is less than 60%), the present invention can produce high-octane light naphtha products. Since the light naphtha generated in the present invention is separated in time, the aromatics in the raw material and the hydrogenation product can be retained to the maximum extent, and the octane number of the light naphtha product is 0.5 to 1.5 units higher than that of the conventional fixed bed reactor process.

[0071] 2. The light naphtha of the present invention has higher selectivity. The present invention uses the separator or distillation chamber to produce the target product light naphtha immediately after hydrogenation, and the target product will not be oversaturated or cracked. The process parameters of the separation section are optimized to separate the target product of the required fraction, thereby increasing the reaction rate and greatly improving the selectivity of light naphtha with high aromatic content.

[0072] 3. In the present invention, the catalyst life is longer. In the present invention, the hydrogenation reaction byproducts such as hydrogen sulfide, ammonia, and small molecule olefins have a very short residence time in the reaction bed, and the reaction heat is also removed from the bed in time. Therefore, the bed coking and clogging problems are basically eliminated.

[0073] 4. The present invention couples the reaction and separation for a second time. After the crude oil passes through the reactor of the present invention, light distillate oil and heavy distillate oil are directly obtained. There is no need to set up a special distillation tower for separation, thus saving investment and energy consumption. At the same time, the heat generated by the hydrogenation reaction process is exactly what the distillation process needs, thus saving thermal energy. The distillation process is mostly used at relatively low pressure or vacuum. The distillation process under high pressure hydrogenation conditions of 8MPaG to 16MPaG is generally considered to fail to achieve the expected effect. However, through the design of the present invention, although the light hydrocarbons after cracking under high pressure are in liquid form, under the action of the separation section and the stripping chamber, these light hydrocarbons can still easily separate from the reaction system, thus achieving the expected separation effect.

[0074] 5. When LCO is used as raw material in the present invention, by adding aromatic extraction facilities, a BTX product with a higher yield can be obtained in the end; by adding a supplementary refining unit, the olefins in the raffinate oil from the aromatic extraction can be fully saturated to obtain an ethylene cracking material with a higher quality. BRIEF DESCRIPTION OF THE DRAWINGS

[0075] Figure 1 and Figure 2 The present invention shows two process flow diagrams for hydrogenating inferior diesel to produce high-octane gasoline or BTX.

[0076] Figure 3 It is a schematic diagram of the structure of a hydrogenation pretreatment reactor.

[0077] Figure 4 It is a schematic diagram of the structure of the hydrogenation reforming reactor.

[0078] Figure 5 Schematic diagram of an annular partition according to an embodiment of a hydrogenation pretreatment reactor of the present invention.

[0079] Figure 6 Schematic diagram of a liquid distribution plate of a hydrogenation pretreatment reactor according to an embodiment of the present invention.

[0080] Figure 7 Schematic diagram of the overflow ring structure according to the present invention.

[0081] Figure 8 A schematic diagram of an embodiment of a gas and liquid inlet distributor for a hydro-modification reactor according to the present invention.

[0082] Fig. 9 It is a schematic structural diagram of an embodiment of a liquid redistributor for a hydroprocessing reactor according to the present invention.

[0083] Fig.10 It is a schematic diagram of an embodiment of a gas-liquid guide plate of a hydrogenation reforming reactor of the present invention.

[0084] Among them: 1-raw oil inlet; 2-hydrogen; 3-hydrogenation pretreatment reactor; 4-hydrogenation pretreatment hydrogen-rich gas; 5-hydrogenation pretreatment naphtha fraction; 6-hydrogenation pretreatment hydrogen-rich non-condensable gas; 7-hydrogenation pretreatment generated oil; 8-hydrogenation reforming reactor raw material; 9-hydrogenation reforming reactor; 10-hydrogenation reforming hydrogen-rich gas; 11-hydrogenation reforming naphtha fraction; 12-hydrogenation reforming hydrogen-rich non-condensable gas; 13-circulating oil pump; 14-unconverted oil; 15-unconverted circulating oil; 16-naphtha product; 17-aromatic extraction system; 18-raffinate oil; 19-BTX product; 20-supplementary refining reactor; 21-supplementary refined oil; 22-high separation tank; 23-ethylene cracking material; 24-desulfurization system; 25-circulating hydrogen compressor; 26, 35-reaction chamber; 27-separator; 28-oil product inlet; 29-hydrogen inlet; 30-lower reaction zone; 31-bed partition; 32, 53-heavy oil bin; 33-liquid redistributor; 34-mixing section; 35-upper reaction zone; 36-separation section; 37-distillation chamber; 38-liquid feed distributor; 3801-material inlet; 3802-distribution pipe; 3803-hole; 39-catalyst bed; 40-gas guide unit; 4001-first cap; 4002-gas guide hole; 4003-first gas guide pipe ;41-hydrogen distribution plate;42-hydrogen distribution chamber;43-guide groove;44-overflow weir;45-liquid redistribution plate;4501-overflow hole;4502-overflow ring;4503-through hole;4503-serrated portion;4503-guide groove;46-gas guide pipe;47-outer wall of guide groove;48-bed baffle;49-heater;50-pressure control valve;51-stripping chamber;52-gas-liquid guide plate;5201-second cap;5202-air hole;5203-second gas guide pipe. DETAILED DESCRIPTION

[0085] 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.

[0086] 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.

[0087] 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.

[0088] 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.

[0089] like Figure 1 The figure shows a schematic diagram of the hydrogenation process of the present invention for producing high-octane gasoline and BTX from low-quality diesel. Low-quality diesel at 200-350°C is fed into the hydrogenation pretreatment reactor 3 from the crude oil inlet 1, and is separated into light components and heavy components in the reactor. The light components pass upward through the reaction bed to be desulfurized, denitrogenated and impurities are removed. Under the action of the tray separation, the required naphtha fraction 5 is separated. The hydrogen-rich gas at the top of the reactor is cooled and then refluxed into the reactor. The hydrogen-rich non-condensable gas 6 is desulfurized by the desulfurization system 24 and then pressurized by the circulating hydrogen compressor 25 and returned to the reactor inlet. The heavy components pass downward through the refined catalyst bed, in which the crude oil is desulfurized, denitrogenated and impurized, part of the olefins are saturated, and part of the macromolecular hydrocarbons are broken into medium-sized hydrocarbon molecules, to obtain the hydrogenation pretreatment generated oil 7, and enter the hydrogenation reforming reactor 9 together with the unconverted circulating oil 15 as the crude oil. After the crude oil enters the hydrogenation reforming reactor 9, the light components in the crude oil enter the separation part upward, and the required naphtha fraction 11 can be obtained. The naphtha fraction 11 obtained from the hydro-modification reactor and the naphtha fraction 5 obtained from the hydro-pretreatment reactor 3 enter the aromatic extraction system 17 together to obtain the BTX product 19. The non-aromatic components in the naphtha fraction enter the supplementary refining reactor 20 as the raffinate oil 18 for hydrogenation saturation to obtain the supplementary refined oil 21, which is separated in the high separation tank 22 to obtain the ethylene cracking material 23, and the hydrogen-rich gas enters the desulfurization facility 24.

[0090] Furthermore, one embodiment of the hydrogenation pretreatment reactor in the present invention is a special-shaped hydrogenation reactor 3, such as Figure 3As shown. After the raw oil 28 enters the reaction chamber 26, it passes through the gas-liquid distributor, and the distillate oil is dispersed into small droplets. Under the action of hydrogen stripping, it is divided into light components and heavy components. The light components contain small molecular hydrogen sulfide, mercaptans and light hydrocarbons. These light components are quickly carried to the mixing section 34, hydrogenated and saturated in the upper reaction zone 35, and further desulfurized. The light components form distillate oil with a temperature gradient in the separation section 36, and the corresponding distillate oil can be extracted as needed. The heavy components contain thiophene, sulfide, polycyclic aromatic hydrocarbons, large molecular alkanes and other components. Under the action of gravity, they enter the different sub-catalyst bed areas in the catalyst bed 30 downward. Hydrogen 29 enters from the bottom and enters the catalyst bed 11 evenly under the action of the gas distributor and the catalyst bed partition 31. The upper layer of the catalyst bed is placed with a hydrorefining catalyst, and the lower layer is placed with a catalyst with a moderate cracking function. Under the action of hydrogen and catalyst, sulfur and nitrogen in the diesel heavy component generate small molecules such as hydrogen sulfide and ammonia, and some polycyclic aromatic hydrocarbons are broken to generate BTX light monocyclic aromatic hydrocarbons. Large molecular hydrocarbons are cracked to generate small molecular hydrocarbons. Under the hydrogenation, hydrogen sulfide, ammonia, small molecular hydrocarbons and some large molecular hydrocarbons are separated from the lower reaction bed 30 and finally enter the separation section 36 for separation. The refined oil 7 after hydrogenation pretreatment enters the next processing unit as raw oil.

[0091] Further, in one or more exemplary embodiments of the present invention, the reaction chamber 30 is an oblate cylinder, such as Figure 3 As shown, the central axis is longitudinal. The plurality of partitions are coaxial annular partitions 31, such as Figure 5 As shown. Further, in one or more exemplary embodiments of the present invention, a plurality of circular holes are distributed on each partition. Further, in one or more exemplary embodiments of the present invention, as Figure 3 As shown, multiple baffles 31 can extend upward to the catalyst bed 30, and the opening rate of the baffles at the bottom that are not in contact with the catalyst bed 30 is less than 70%. The lower opening rate is conducive to increasing resistance, so that hydrogen can enter the catalyst bed 30 upward as much as possible, further playing the role of a gas distributor. The opening rate of the baffles in the catalyst bed 11 is greater than 50%, which is conducive to more effective use of the catalyst.

[0092] Further, in one or more exemplary embodiments of the present invention, the ratio of the cross-sectional area of ​​the separator 27 to that of the reaction chamber 26 is 1:1.2 to 1:10.

[0093] Further, in one or more exemplary embodiments of the present invention, the separator 27 includes, from bottom to top, a mixing section 34 , an upper reaction zone 35 , and a separation section 36 .

[0094] Further, in one or more exemplary embodiments of the present invention, the liquid distribution assembly 33 between the separator 27 and the reaction chamber 26 includes a liquid distribution plate and a distribution cone (such as Figure 6 ). The liquid distribution plate is arranged above the liquid distributor, and the liquid distribution plate has the same shape as the top surface of the porous catalyst layer. A plurality of holes are evenly provided on the liquid distribution plate, and an overflow ring 4502 is arranged around the holes. An overflow portion is arranged on the outer edge of the distribution plate (not shown in the figure). The distribution cone is arranged at the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of holes, and a similar overflow ring is arranged around the holes (not shown in the figure).

[0095] Further, in one or more exemplary embodiments of the present invention, the area of ​​the liquid distribution plate is 10% to 100% of the cross-section of the catalyst bed, preferably 60% to 100%. The opening rate of the liquid distribution plate is 5% to 90%, the diameter of the hole is 5 to 100 mm, the height of the overflow ring is 1 to 30 mm, and the height of the overflow portion of the edge of the liquid distribution plate can be the same as the height of the overflow ring, but the present invention is not limited thereto. The top angle of the distribution cone is greater than 90°, and the flatter the top angle of the distribution cone, the better, so that the flow rate of the droplets in the liquid distribution plate can be relatively flat. The opening rate of the distribution cone is 5% to 80%, and the bottom area of ​​the distribution cone is 2% to 15% of the area of ​​the liquid distribution plate.

[0096] Preferably, but not limiting, in one or more exemplary embodiments of the present invention, as shown in FIG. Figure 7 As shown, a serration 4504 is provided on the inner side of the overflow ring, the serration is bent downward, and a guide groove 4505 is provided on the serration. Exemplarily, the guide groove is opened along the center of the serration.

[0097] Further, in one or more exemplary embodiments of the present invention, the void fraction of the catalyst bed is 15% to 85%, preferably 20% to 75%.

[0098] Furthermore, one of the embodiments of the hydrogenation reforming reactor in the present invention is a hydrogenation distillation reactor 9. The hydrogenation distillation reactor of the present invention includes a reaction chamber 35, a distillation chamber 37, a stripping chamber 51 and a heavy oil bin 53. Among them, a catalyst bed arranged in an annular manner is provided in the reaction chamber 35, and the catalyst bed can be provided in one or more layers. After the raw oil 8 enters the reaction chamber 35, it is dispersed into fine droplets under the action of the liquid distributor 38, wherein the light components containing high aromatics are not in contact with the catalyst bed 39 and are brought to the distillation chamber 37 for separation under the action of hydrogen. The hole diameter of the liquid phase distributor 38 is 1 mm to 20 mm. Other heavy components enter the catalyst bed 39 downward. The hydrogen 2-2 fills the hydrogen distribution plate in the hydrogen distribution chamber and enters the catalyst bed 39 upward uniformly, and contacts the downward heavy fraction oil in countercurrent, and reacts under the process conditions of the present invention. The nitrogen-containing compounds in the heavy fraction raw oil are hydrogenated to generate intermediate products, such as indole nitrogen heterocycle hydrogenation to generate dihydroindole, carbazole hydrogenation to generate tetrahydrocarbazole, pyridine to generate piperidine, etc., polycyclic aromatic hydrocarbons, long-chain alkylbenzenes and other parts are partially ring-opened and broken to generate small molecular hydrocarbons or light aromatics, and sulfur-containing compounds such as mercaptans and thioethers generate hydrogen sulfide. These small molecules are quickly transferred to the bed by hydrogen, and the reaction moves to the right. The heavy fraction intermediate products after the reaction of the first bed continue to move to the second bed, and the middle fraction oil in the second bed continues to be hydrogenated, and the generated light aromatic hydrocarbons, small molecular hydrocarbons, ammonia, and thiophene sulfides are hydrogenated to hydrogen sulfide. The light components generated by hydrogenation in the second bed are also quickly brought to the stripping chamber 51 by hydrogen, and the rest of the beds are analogous. In this way, each bed in the reaction chamber continues to hydrogenate the heavy fraction that is difficult to hydrogenate at a slow rate, thereby increasing the concentration of reactants and reducing the concentration of products, thereby accelerating the reaction rate and selectivity. During the stripping process, a small amount of heavy components in the stripping chamber 51 will aggregate. When the aggregated liquid is larger, it falls into the gas-liquid guide plate 52 in the stripping chamber. Fig.10 As shown in the left figure, the gas is finally diverted to the guide groove 43 and enters the lower catalyst. The gas at the bottom of the guide plate 52 passes through the air hole 5202 in the second air guide pipe 5203, and the second cap can prevent the downward liquid from passing through the air hole 5202. Preferably, but not restrictively, the gas-liquid guide plate 52 is conical, with a top angle greater than 60° and less than 180°, and generally greater than 90°, to increase the gas phase channel space. The opening rate of the gas-liquid guide plate 52 is 5% to 99%, the diameter of the air hole 5202 can be 3mm to 100mm, and the height of the second air guide pipe 5203 is 1mm to 30mm.

[0099] Further, in the present invention, the feedstock oil passes through the bed downwards through the overflow weir 44 and enters the liquid redistribution plate 45 under the guidance of the guide groove 43. Hydrogen is discharged from the first gas guide pipe 4003 ( Fig.10) in the gas guide hole 4002 upwardly enters the catalyst bed layer, and a cap 4001 is provided on the upper part of the guide pipe to prevent the raw oil in the bed layer from entering the gas guide pipe. Preferably, but not restrictively, the opening rate of the liquid redistribution plate is 5% to 90%, the diameter of the hole 4503 is 3 mm to 100 mm, the height of the first gas guide pipe 4003 is 20 mm to 500 mm, and the height difference between the overflow weir 44 and the first gas guide pipe 4003 is -50 to -200 mm. The first cap 4001 is kept directly above the first gas guide pipe 802.

[0100] The catalyst used in the present invention is a conventional hydrorefining catalyst and a hydrocracking catalyst in the art. In order to better achieve the effect of the present invention, the present invention provides a corresponding method for preparing the hydrorefining catalyst and the hydrocracking catalyst, but it should not be understood as a limitation on the catalyst used in the present invention.

[0101] The hydrotreating catalyst of the present invention is preferably a hydrotreating catalyst with high desulfurization and denitrification activity and less 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 silicon aluminum, 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.

[0102] The preferred catalyst active component of the hydrocracking catalyst of the present invention is 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 component of the catalyst includes one or more of alumina, silicon-containing alumina and molecular sieves, preferably a molecular sieve, and the molecular sieve may be a Y-type molecular sieve. In a preferred embodiment, based on the weight of the catalyst, it contains 10wt% to 35wt% of the metal of Group VIB in terms of oxide, 3wt% to 15wt% of the metal of Group VIII in terms of oxide, 5wt% to 40wt% of the molecular sieve, 15wt% to 72wt% of amorphous silicon aluminum, and 10wt% to 67wt% of alumina; the specific surface area of ​​the catalyst is 100 to 650m 2 / g, and the pore volume is 0.15~0.50mL / g.

[0103] Preferably, the hydrotreating catalyst and hydrocracking catalyst of the present invention are preferably porous catalysts, shaped catalysts and / or honeycomb 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 pore diameter or pore side length of the honeycomb catalyst is 1 to 50 mm, preferably 3 to 15 mm; the void ratio of the catalyst bed is recommended to be 15% to 85%, preferably 20% to 75%.

[0104] Furthermore, the catalyst used in the present invention is in the shape of a honeycomb, ball ring, Raschig ring, rectangular saddle ring, saddle, open-hole ring type, half ring, step ring, double arc, Haier ring, conjugated ring, flat ring, wreath, hollow sphere and other fillers or other porous shapes.

[0105] The following examples further illustrate a method for hydrogenating low-quality diesel oil that can improve the selectivity and yield of naphtha 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.

[0106] The experimental methods in the following examples, unless otherwise specified, are conventional methods in the art.

[0107] Example 1

[0108] Using the present invention Figure 1 The hot crude oil enters the reactor from the middle of the hydrogenation pretreatment reactor. Figure 3The shaped reactor is shown. The catalyst bed in the lower reaction zone is filled with catalyst A and catalyst B from top to bottom, and the volume ratio of the two catalysts is VA:VB=2:1. Catalyst C is placed in the upper reaction zone. The catalyst properties are shown in Table 2. The diameter-to-height ratio of the catalyst bed in the reaction chamber is 4:1. An annular partition is arranged in the catalyst bed, and the number of partitions is 5. There are multiple holes distributed on the partition; the partition extends upward to the catalyst layer, the opening rate of the partition below the catalyst layer is 40%, and the opening rate of the partition in the catalyst layer is 70%. After the crude oil enters the hydrogenation pretreatment reaction chamber, the light components in the crude oil, such as light aromatics BTX, small molecular alkanes, hydrogen sulfide, mercaptans, etc., move upward into the upper reaction bed under the hydrogenation, and are desulfurized and denitrogenated. The light component distillate oil containing high aromatics is separated in the separation section. The heavy components in the feedstock oil, such as polycyclic aromatic hydrocarbons such as C8 and C10, alkylbenzenes, and long-chain alkanes, enter the lower reaction bed for hydrorefining, hydrocracking and ring-opening reactions. The small molecular components such as hydrogen sulfide, ammonia, and light aromatic BTX produced immediately leave the reaction bed under the action of hydrogen and move to the separation section, which accelerates the hydrogenation reaction of the heavy hydrocarbons to move to the right. The heavy components after hydrogenation continue to move downward and enter the hydro-reforming reactor. The diameter ratio of the distillation chamber to the reaction chamber of the hydro-reforming reactor is 1:3, and the diameter-to-height ratio of the catalyst bed in the reaction chamber is 5:1. The heavy components in the hydro-reforming reactor continue to undergo hydrocracking and ring-opening reactions in the reaction bed. The light BTX and small molecular hydrocarbons produced by the reaction enter the stripping chamber through the gap between the stripping chamber and the bed. Finally, the light distillate oil containing high aromatics is further separated through the distillation chamber and mixed with the high aromatic naphtha fraction extracted from the hydro-pretreatment reactor as the raw material for downstream processing. The remaining heavy components enter the liquid redistributor through the overflow weir, and then enter the next catalyst bed. The light components produced in the next catalyst bed enter the stripping chamber. After passing through 15 beds in sequence, the unconverted oil enters the heavy oil tank and then returns to the entrance of the hydrogenation reforming reaction chamber. The process conditions are shown in Table 3.

[0109] In the embodiment, the area of ​​the liquid distribution plate of the pretreatment reaction chamber is 70% of the cross-section of the catalyst bed. The opening rate of the liquid distribution plate and the distribution cone is 50%, the diameter of the hole is 10mm, and the height of the overflow ring is 10mm. The top angle of the distribution cone is 120°, and the bottom area of ​​the distribution cone is 10% of the area of ​​the liquid distribution plate. In the hydrogenation reforming reactor, the hole diameter of the liquid feed distributor 38 is 4mm. The first air guide pipe 4003 on the gas guide unit 40 is 50mm high, and the height of the overflow weir 44 is 30mm. The diameter of the air hole 520 in the conical gas-liquid guide plate 52 is 30mm, and the height of the second air guide pipe 5203 is 20mm. The opening rate of the liquid redistribution plate 45 is 85%, the diameter of the through hole 4503 is 20mm, and the height is 15mm. The distribution area of ​​the hydrogen distribution plate 41 and the liquid feed distributor 38 accounts for 95% of the bed distribution area.

[0110] The properties of the raw materials of this example are shown in Table 1. The catalysts are shown in Table 2, the process operating conditions are shown in Table 3, and the product distribution of the examples and comparative examples is shown in Table 4. The simulated temperature rise change of the catalyst layer is shown in Table 5.

[0111] Example 2

[0112] The difference between this embodiment and embodiment 1 is that the raw oil used is 2# coking diesel. The other conditions are the same as those in embodiment 1.

[0113] Example 3

[0114] The difference between this embodiment and embodiment 1 is that the raw oil used is 3# straight-run diesel. The other conditions are the same as those in embodiment 1.

[0115] Example 4

[0116] The difference between this embodiment and embodiment 1 is that the number of beds in the hydro-reforming reactor is 20. The other conditions are the same as those in embodiment 1.

[0117] Example 5

[0118] The difference between this embodiment and embodiment 1 is that the volume ratio of catalyst A to catalyst B in the hydrogenation pretreatment reactor is 5:1. The other conditions are the same as those in embodiment 1.

[0119] Example 6

[0120] The difference between this embodiment and embodiment 1 is that the hydrogenation reforming catalyst is prepared using a conventional carrier, and the diameter of the prepared catalyst E is 1.8 mm, the length is 5-8 mm, and the bed void ratio is 30%. The other conditions are the same as those in embodiment 1.

[0121] Example 7

[0122] The same as Example 1, except that the diameter-to-height ratio of the catalyst bed in the special-shaped hydrogenation reactor is 1:3.

[0123] Example 8

[0124] The same as Example 1, except that the light distillate oil extracted from the hydropretreatment reactor and the light distillate oil extracted from the hydroupgrading reactor enter the aromatics extraction device together, and the raffinate oil enters the supplementary refining reactor for treatment.

[0125] The catalyst properties of the above-mentioned embodiments and comparative examples are shown in Table 1, the raw oil properties 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. The temperature rise of the simulated hydrocracking bed of the embodiments and comparative examples is shown in Table 5.

[0126] Comparative Example 1

[0127] A conventional two-stage hydrogenation process, i.e., 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 in a separation tower and sent out as a product. In the comparative example, the refining reactor is filled with catalyst A and catalyst B in a volume ratio of 2:1, and the cracking reactor is filled with catalyst D. The tail oil full circulation process is adopted, and the tail oil is intermittently thrown out. The remaining process conditions are the same as those in Example 1.

[0128] Table 1 Properties of crude oil

[0129]

[0130] Table 2 Catalyst properties

[0131]

[0132] Table 3 Process conditions of the embodiments and comparative examples

[0133]

[0134] Table 4 Examples and product yields

[0135]

[0136] From the results in Table 4, it can be seen that by adopting the method of the present invention, the yield of naphtha and light aromatic BTX in naphtha is significantly higher than that of Comparative Example 1.

[0137] The laboratory simulated and calculated the bed reaction temperature distribution of the embodiment and the comparative example by using ansys 19.0 version software. The simulation conditions were input according to the actual data of the embodiment and the comparative example. The simulation results show that the center temperature of the traditional fixed bed is the highest, and the temperature change is normally distributed from the inlet end to the outlet end. The bed temperature of the reactor of the present invention is relatively uniform. The simulated temperature rise change of the bed is shown in Table 5.

[0138] Table 5 Simulated temperature rise changes of the 1# diesel bed during treatment

[0139]

[0140] It can be seen from the results in Table 5 that the temperature difference of the catalyst bed in Example 1 of the present invention is significantly lower than that in Comparative Example 1. The temperature difference of the catalyst bed in the hydrogenation pretreatment and hydrogenation reforming reactor of Example 1 of the present invention is significantly lower than that in Comparative Example 1, and the maximum radial temperature difference is reduced from 22.3°C of the traditional fixed bed to 0.5°C. The difference between the average temperature of the embodiment and the controlled temperature is only 1.3°C to 1.6°C, and the difference between the average temperature of the comparative example and the controlled temperature reaches 9.7°C to 10.4°C, indicating that the reaction heat of the reactor of the present invention is taken out of the catalyst bed in time, without accumulation in each bed, and the operation stability and controllability are higher.

Claims

1. A hydroprocessing device for producing more BTX from inferior diesel, characterized in that: The device comprises: The hydrogenation pretreatment reactor is a special-shaped hydrogenation reactor, comprising a heavy oil tank, a reaction chamber and a separator which are sequentially connected from bottom to top, and the reaction chamber is sequentially provided with: a liquid distribution component, a catalyst bed and a hydrogen distribution chamber from top to bottom; The liquid distribution assembly of the hydrogenation pretreatment reactor includes a liquid distributor, and a liquid distribution plate and a distribution cone arranged above the liquid distributor; the separator includes a mixing section, a separation section and a stabilization section from bottom to top; The hydrogenation pretreatment reactor comprises at least one catalyst bed, and the ratio of the equivalent diameter of the cross-sectional area of ​​the catalyst bed to the total height of the catalyst bed is 2:1 to 10:1; The oil generated by the hydropretreatment enters the liquid phase feed distributor of the hydroupgrading reactor; The hydro-reforming reactor is a hydro-distillation reactor, comprising: a reaction chamber, in which a catalyst bed layer arranged in an annular manner is arranged, and the liquid raw oil and the hydrogen feed are in reverse contact in the catalyst bed layer; a stripping chamber, which is located in the middle of the hydro-distillation reactor and is a gas phase channel that runs through the reactor from top to bottom, and the gas phase product after the hydrogenation reaction of each catalyst bed layer enters the gas phase channel under the action of stripping; a distillation chamber, which is arranged at the upper part of the hydro-distillation reactor, and the gas phase product is distilled and separated in the distillation chamber; The liquid phase feedstock oil is a mixed component of the oil produced by hydrogenation pretreatment and the unreacted distillate oil in the hydrogenation reforming reactor; The reaction chamber of the hydrodistillation reactor is provided with: a liquid phase feed distributor, which is arranged above the uppermost catalyst bed layer and is an annular tube type or an annular belt type; a hydrogen feed distributor, which is arranged at the lower part of each catalyst bed layer and is an annular tube type or an annular belt type; The diameter-to-height ratio of each catalyst bed in the hydro-modification reactor is 2:1 to 20:

1.

2. The hydroprocessing device according to claim 1, characterized in that: The cross section of the catalyst bed in the reaction chamber of the hydrogenation pretreatment reactor is rectangular or circular, and the bed height is 100 to 5000 mm.

3. The hydroprocessing device according to claim 2, characterized in that: A plurality of partitions are arranged in parallel in the vertical direction in the hydrogen distribution chamber and the catalyst bed in the reaction chamber of the hydrogenation pretreatment reactor. The plurality of partitions divide the hydrogen distribution chamber into a plurality of air inlet units. Each air inlet unit has at least one hydrogen inlet at the bottom. Each partition has a plurality of holes. The partition extends upward to the catalyst layer.

4. The hydroprocessing device according to claim 3, characterized in that: The opening rate of the partition below the catalyst bed of the hydrogenation pretreatment reactor is less than 70%, and the opening rate of the partition in the catalyst layer is greater than 50%.

5. The hydroprocessing device according to claim 3, characterized in that: In the hydrogenation pretreatment reactor, there are 1 to 3 hydrogen distributors at the bottom of the catalyst bed between every two adjacent partitions, and the distribution area of ​​hydrogen from all hydrogen distributors in the partition area when it reaches the bottom of the regional bed should cover the entire regional bed bottom.

6. The hydroprocessing device according to claim 3, characterized in that: The heavy oil bunker is arranged at the bottom center of the hydrogenation pretreatment reaction chamber, and the heavy oil bunker is connected with the multiple air inlet units.

7. The hydroprocessing device according to claim 1, characterized in that: In the hydrogenation pretreatment reactor, the area of ​​the liquid distribution plate is 10% to 100% of the cross section of the catalyst bed.

8. The hydroprocessing device according to claim 1, characterized in that: In the hydrogenation pretreatment reactor, a plurality of first through holes are evenly arranged on the liquid distribution plate, a first overflow ring is arranged around the first through holes, and an overflow portion is arranged at the outer edge of the liquid distribution plate.

9. The hydroprocessing device according to claim 8, characterized in that: A serration portion is provided on the inner side of the first overflow ring. The serration portion is bent downward and a guide groove is provided on the serration portion.

10. The hydroprocessing device according to claim 8, characterized in that: The distribution cone is arranged at the upper center of the liquid distribution plate, and a plurality of holes are arranged on the distribution cone, and overflow rings are arranged around the holes.

11. The hydroprocessing device according to claim 1, characterized in that: The hydrogen feed pipe at the bottom of the lower reaction section of the reaction chamber of the hydrogenation pretreatment reactor is provided with multiple inlets, and each hydrogen feed pipe corresponds to the catalyst bed area between two partitions, so that after hydrogen comes out of each liquid distributor, it can pass upward through the top reaction zone.

12. The hydroprocessing device according to claim 1, characterized in that: The hydrogen feed distributor is arranged in a relatively closed space.

13. The hydroprocessing device according to claim 1, characterized in that: A gas guide unit is provided between the hydrogen feed distributor and the bottom of the catalyst bed in the reaction chamber of the hydrogenation reforming reactor; the gas guide unit comprises: a guide plate, which is provided at the bottom of the catalyst bed and is used to receive the liquid phase product after the hydrogenation reaction of the catalyst bed; a guide pipe, which is evenly arranged on the guide plate and guides the hydrogen feed upward to the catalyst bed, and a first cap is provided on the top of the guide pipe; an overflow weir, which is close to one side of the stripping chamber and is used to overflow the liquid phase product on the guide plate to the next catalyst bed.

14. The hydroprocessing device according to claim 13, characterized in that: An annular liquid redistribution plate with evenly-spaced through holes is provided below the overflow weir.

15. The hydroprocessing device according to claim 14, characterized in that: The liquid redistribution plate comprises: an overflow ring arranged around the top of the through hole for uniformly guiding the liquid phase product to the next catalyst bed; a baffle arranged on the side away from the stripping chamber, the upper edge of the baffle being higher than the overflow ring.

16. The hydroprocessing device according to claim 15, characterized in that: The overflow ring comprises: a sawtooth portion, which is bent downward from the outer edge of the through hole toward the center; and a drainage groove, which is arranged on the sawtooth portion and opened along the center of the sawtooth portion.

17. The hydroprocessing device according to claim 1, characterized in that: The stripping chamber of the hydrogenation reforming reactor is provided with a conical gas-liquid guide plate, and the number and position of the gas-liquid guide plate are matched with the catalyst bed.

18. The hydroprocessing device according to claim 17, characterized in that: The conical surface of the gas-liquid guide plate is provided with air holes, and the top of the air holes is provided with a second cap; the gas phase products after the hydrogenation reaction rise from the air holes and converge into the distillation chamber, and part of the heavy component liquid phase products drip onto the gas-liquid guide plate and flow back to the reaction chamber.

19. The hydroprocessing device according to claim 1, characterized in that: The distillation chamber of the hydrogenation reforming reactor comprises a mixing section and a separation section, and a packing or a tower plate is arranged in the separation section.

20. The hydroprocessing device according to claim 1, characterized in that In the hydrogenation pretreatment reactor and the hydrogenation reforming reactor, the ratio of the diameter of the upper separator or distillation chamber to the diameter of the lower reaction chamber is 1:1.2 to 1:

50.

21. The hydroprocessing device according to claim 1, characterized in that: The reaction chamber of the hydrogenation pretreatment reactor is a horizontal storage tank, the axial direction of which is arranged in the transverse direction, and ends are provided at both ends of the horizontal storage tank; or the reaction chamber is a flat cylindrical tank, the axial direction of which is arranged in the longitudinal direction.

22. A hydroprocessing method for producing more BTX from inferior diesel, using the hydroprocessing device according to any one of claims 1 to 21, characterized in that: The steps include: (1) After the low-quality diesel fuel enters the liquid distribution assembly of the hydrotreating pretreatment reactor and is dispersed, the lighter part is carried upward by the hydrogen flowing upward from the bottom, and the heavier part enters the hydrotreating pretreatment bed layer downward to generate hydrotreating pretreatment product oil; (2) The light components described in step (1) are separated upward in the separation section of the hydrogenation pretreatment reactor, and the naphtha product is extracted from the side line; (3) The hydropretreatment product oil obtained in step (1) enters the liquid phase feed distributor of the hydroprocessing reactor, and the light components in the distillate oil dispersed by the liquid phase feed distributor enter the separation section upward; the heavy components enter the multi-bed hydroprocessing reaction zone downward; (4) Every time the heavy component in step (3) enters a catalyst bed, the light component produced enters the separation section in step (3) in a timely manner; and finally, the unreacted distillate oil is fully circulated and enters the inlet of the hydro-reforming reactor; (5) After the light components in step (3) are separated in the separation section, a naphtha fraction is extracted from the side line and mixed with the naphtha fraction in step (2).

23. The hydroprocessing method according to claim 22, characterized in that: The method further comprises step (6): the naphtha fraction product described in step (5) enters an aromatic extraction unit to separate light aromatic BTX product and light non-aromatic fraction oil.

24. The hydroprocessing method according to claim 22, characterized in that The operating conditions of the hydrogenation pretreatment reactor in step (1) are as follows: the average reaction temperature is 200-450°C, the reaction pressure is 3-20 MPa, the tower top reflux ratio is 1.2-4.5, the hydrogen-to-oil volume ratio is 100:1-3000:1; the raw material inferior diesel flow rate and the hydrogenation pretreatment reactor catalyst volume ratio is 0.1-10.0h -1 .

25. The hydroprocessing method according to claim 22, characterized in that: The separation section described in step (2) is provided with a refining reaction bed layer for removing sulfides from light aromatic components.

26. The hydroprocessing method according to claim 22, characterized in that: The operating conditions of the separation section in step (2) are as follows: the reflux temperature at the top of the separation section is 50°C to 70°C, the reflux ratio is 1.1 to 10.0, the side line extraction temperature is 150°C to 185°C, and the pressure of the separation section is 3 to 20 MPa.

27. The hydroprocessing method according to claim 22, characterized in that: In step (1), the hydrogenation pretreatment reaction zone is divided into a plurality of sub-reaction zones along the longitudinal direction, so that when the hydrogen rises to the bottom of the bed through the distributor, the diffusion area of ​​the hydrogen can cover the cross-sectional area of ​​the bed at the bottom of the sub-reaction zone, and each sub-reaction zone is provided with a separate hydrogen inlet.

28. The hydroprocessing method according to claim 22, characterized in that: Two catalysts are placed in the hydrotreating reaction zone of step (1); a hydrotreating catalyst is placed in the upper part, and a hydrocracking catalyst is placed in the lower part; the volume ratio of the hydrotreating catalyst to the hydrocracking catalyst is 1:1 to 10:

1.

29. The hydroprocessing method according to claim 22, characterized in that The operating conditions of the hydrogenation reforming reaction zone in step (3) are as follows: the average operating temperature is 230°C to 450°C, the operating pressure is 3MPa to 20MPa, and the hydrogen-to-oil volume ratio is 200:1 to 4000:

1.

30. The hydroprocessing method according to claim 22, characterized in that: The operating conditions of the separation section in the hydrogenation reactor in step (3) are as follows: the reflux temperature at the top of the separation section is 50°C to 70°C, the reflux ratio is 1.1 to 10.0, the side line extraction temperature is 150°C to 185°C; and the pressure of the separation section is 3 to 20 MPa.

31. The hydroprocessing method according to claim 22, characterized in that: The number of reaction zone beds in the reaction section of the hydrogenation reactor in step (3) is 1 to 30, and the height of each bed is 50 to 500 mm.

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