Hydrotreatment method and device for catalyzing diesel oil to produce BTX in high yield

Abstract

The invention discloses a treatment method for increasing BTX yield by hydrogenation of LCO diesel oil. By adopting a multi-bed hydrogenation rectification reactor and a self-circulation system, the purposes of echelon reaction of raw oil, immediate production and immediate removal, effective control of the reaction depth and improvement of the yield of BTX products are achieved, meanwhile, the risk of temperature runaway or flooding of beds can be eliminated, and the service life of a catalyst is prolonged.

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, and then extracts the light fractions to obtain light aromatic fractions and light non-aromatic fractions; the heavy fractions and light non-aromatic fractions are successively subjected to hydrofining and hydromodification reactors to finally obtain gasoline components with 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, with high energy consumption and high investment. In addition, the hydrocracking reactor and hydrofining reactor use traditional fixed-bed reactors, and there is a situation of excessive hydrogenation or cracking of the distillate oil, which ultimately affects the yield of light aromatics.

[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. Summary of the invention

[0011] In view of the shortcomings of the prior art, the purpose of the present invention is to provide a method and device for producing more BTX by hydrogenating LCO diesel. By adopting a multi-bed hydrogenation distillation reactor and a self-circulating system, the raw oil can be reacted in stages, and the reaction depth can be effectively controlled to achieve the purpose of increasing the yield of BTX products. At the same time, the risk of "overheating" or flooding of the bed can be eliminated, and the service life of the catalyst can be extended.

[0012] The first aspect of the present invention provides a method for hydrotreating LCO diesel to produce more BTX, comprising the following steps: (1) The raw materials LCO diesel and hydrogen enter the hydrogenation pretreatment reaction zone for hydrogenation treatment and are separated into gas and liquid; (2) The liquid produced in step (1) enters the lower bed of the hydrocracking reactor, and after being dispersed by the liquid distribution assembly, the lighter portion is carried to the middle of the reactor by the hydrogen flowing upward from the bottom, and the heavier portion enters the hydrocracking reaction zone of the multiple beds downward to contact with the hydrogen in countercurrent to carry out the hydrocracking reaction; the light products after hydrocracking in each bed of the lower hydrocracking reaction zone are separated from the reaction bed in time, and the unconverted oil enters the next bed for reaction, and the raw oil of each bed is the unconverted oil of the previous bed; the light products separated from each bed are gathered together and enter the middle of the reactor; (3) A stripping chamber is provided in the middle of the hydrocracking reaction zone, and a distillation chamber is provided in the upper part. The light components described in steps (2) and (3) first enter the stripping chamber in the middle for stripping, and then enter the distillation chamber in the upper part of the hydrocracking reactor for separation. A naphtha product rich in BTX is extracted from the side product line; (4) After the hydrogenation reaction occurs in the reactor bed in step (2), the light components produced enter the upper distillation chamber of the reactor for distillation separation, and the heavy monocyclic aromatic hydrocarbon-rich fraction extracted from the side return line enters the upper reaction zone of the hydrocracking reactor for hydrogenation reaction; the light components produced by the reaction enter the middle part of the reactor, and the heavy components enter the lower hydrocracking reaction zone; (5) The uncracked heavy components in the hydrocracking reaction zone are at least partially circulated into the hydrocracking reaction zone by a pump for reaction.

[0013] Furthermore, the BTX-rich naphtha fraction product described in the above step (4) enters an aromatic extraction unit to separate light aromatic BTX products and light non-aromatic fraction oil.

[0014] Furthermore, the above light non-aromatic fraction oil enters into a supplementary hydrotreating reactor, and the refined oil is used as ethylene cracking feed.

[0015] Furthermore, in the above technical solution, the hydroprocessing reaction zone described in step (1) is a fixed bed hydroprocessing reactor, which contains a hydroprocessing catalyst bed. The hydroprocessing catalyst bed can be entirely hydrorefining catalyst, or a mixed loading of refining catalyst and other hydroprocessing catalysts. Preferably, the hydrorefining catalyst is placed on the upper part and the hydrocracking catalyst is placed on the lower part. The volume ratio of the hydrorefining catalyst to the hydrocracking catalyst is 1:1 to 10:1, preferably 3:1 to 5:1. By using two different types of catalysts, the feedstock oil can be desulfurized and denitrogenated, and after appropriate cracking, the feedstock oil can be converted into a clean medium molecular distillate as a feedstock for subsequent hydrocracking reactions.

[0016] 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.

[0017] Furthermore, in the above technical solution, the hydrocracking reaction zone in step (2) comprises at least two catalyst beds, 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.

[0018] Furthermore, in the above technical solution, the catalyst bed in the hydrocracking reaction zone of step (2) is divided into at least two parts, an upper part and an lower part, each part has a liquid feed inlet, and a hydrogen feed inlet is arranged at the bottom of each bed in each part.

[0019] Furthermore, in the above technical solution, the temperature of the side line location for extracting BTX-rich naphtha in step (3) is 140°C to 170°C, preferably 145°C to 160°C.

[0020] Furthermore, in the above technical solution, a reflux line is provided at the top of the distillation chamber in step (3), and the reflux temperature is 50°C to 70°C, preferably 55°C to 65°C.

[0021] Furthermore, in the above technical solution, the temperature of the side line position for extracting the heavy monocyclic aromatic hydrocarbon fraction in step (4) is 160°C to 250°C, preferably 165°C to 220°C.

[0022] Furthermore, in the above technical solution, the feedstock oil of the present invention is catalytic cracking light cycle oil (LCO). The total aromatic content is as high as 50% to 90%, 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 to 1.5 wt%, and the nitrogen content is 100 to 750 ppm.

[0023] 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 hydrogen-oil volume ratio is 100:1-3000:1. The feedstock 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, and hydrogen to oil volume ratio of 200:1 to 1500:1.

[0024] Furthermore, in the above technical solution, the average reaction temperature of the upper reaction zone of the hydrocracking reactor described in step (4) is 210-430°C, the reaction pressure is 3-20MPa, and the volume ratio of hydrogen to oil is 100:1-3500:1. The volume ratio of the feedstock oil flow rate to the reactor catalyst is 0.1-10.0h -1 The preferred operating conditions are: average reaction temperature of 250°C to 390°C, reaction pressure of 4 to 18 MPa, and hydrogen to oil volume ratio of 200:1 to 1500:1.

[0025] Furthermore, in the above technical solution, the average reaction temperature of the lower reaction zone of the hydrocracking reactor described in step (4) is 230-430°C, the reaction pressure is 3-20MPa, and the volume ratio of hydrogen to oil is 100:1-4000:1. The volume ratio of the feedstock oil flow rate to the reactor catalyst is 0.1-10.0h -1 The preferred operating conditions are: average reaction temperature of 280°C to 420°C, reaction pressure of 4 to 18 MPa, and hydrogen to oil volume ratio of 200:1 to 1800:1.

[0026] Furthermore, the operating conditions of the distillation chamber in step (3) and step (4) are as follows: the reflux temperature at the top of the distillation chamber 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 product extraction temperature is 150°C to 185°C, preferably 160°C to 170°C. The side return line extraction temperature is 150°C to 185°C, preferably 160°C to 170°C. The pressure of the distillation chamber is 3 to 20 MPa, preferably 4 to 18 MPa.

[0027] Furthermore, the number of reaction zone beds in the reaction section of the hydrocracking 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.

[0028] 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.

[0029] The second aspect of the present invention also provides a LCO diesel hydroprocessing device for producing more BTX, the device comprising: (1) Hydrogenation pretreatment reactor and hydrotreatment reactor are used to hydrotreat the low-quality diesel raw material; (2) The hydrocracking reactor is a multi-bed hydrodistillation reactor, comprising: A reaction chamber is provided with a catalyst bed arranged in an annular manner, and the liquid raw oil and the hydrogen feed are in reverse contact in the catalyst bed; The stripping chamber is located in the middle of the hydrodistillation reactor and is a gas phase channel that runs through the reactor from top to bottom. The gas phase product after the hydrogenation reaction of each catalyst bed enters the gas phase channel under the action of stripping; The distillation chamber is arranged at the upper part of the hydrogenation distillation reactor, and the gaseous product is distilled and separated in the distillation chamber, and a return line is arranged in the middle part.

[0030] Furthermore, in the above technical solution, 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 0.2:1 to 10:1, preferably 2:1 to 6:1. The reactor bed described in the present invention is preferably of equal diameter, that is, the cross-sectional area at different positions within the entire catalyst bed is the same. The cross-sectional area of ​​the catalyst bed is usually the same as the cross-sectional area of ​​the reaction chamber in the reactor, and the cross section refers to the top view cross section, that is, the cross section perpendicular to the vertical line in the reaction chamber. If there are differences in the cross-sectional area of ​​the reactor within the catalyst bed height range, the cross-sectional area here refers to the average value of the cross-sectional area of ​​the catalyst bed or the cross-sectional area of ​​the reaction chamber at various locations within the catalyst bed range.

[0031] Furthermore, in the above technical solution, the reaction chamber is preferably a cylindrical storage tank, the axial direction of which is arranged along the longitudinal direction, or a horizontal storage tank, the axial direction of which is arranged along the transverse direction, and the two ends of the horizontal storage tank are provided with heads.

[0032] The catalyst bed 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 hydrogen sulfide and ammonia generated in each sub-bed 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 the case in traditional reactors, will not occur.

[0033] Furthermore, in the above technical solution, a conical gas-liquid guide plate may be provided in the stripping chamber, and the number and position of the gas-liquid guide plate match the catalyst bed. A through hole is provided on the gas-liquid guide plate, a cofferdam is provided around the through hole, and a cap is preferably provided above the through hole. The opening rate of the gas-liquid guide plate is 5% to 90%, the diameter of the through hole is 3 to 100 mm, and the height of the cofferdam is 1 to 30 mm. The gas phase products after the hydrogenation reaction rise from the pores 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.

[0034] Furthermore, in the above technical solution, 2 to 3 liquid feed distributors are arranged in the reaction chamber, which are arranged above different catalyst beds in the upper part of the reaction zone, preferably in the form of an annular tube or an annular belt; hydrogen feed distributors are arranged at the bottom of each catalyst bed, preferably in the form of an annular tube or an annular belt. However, the present invention is not limited to the preferred gas and liquid feed distributors.

[0035] Further, in the above technical solution, a gas guide unit may be provided between the hydrogen feed distributor and the bottom of the catalyst bed, 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 cap is provided on the top of the guide pipe; an overflow weir, which is close to one side of the distillation chamber, is used to overflow the liquid phase product on the guide plate to the next catalyst bed. Preferably, but not restrictively, the opening rate of the guide plate is 5% to 90%, the diameter of the guide hole is 3mm to 100mm, the height of the guide pipe is 10mm to 300mm, and the height difference between the overflow weir and the guide pipe is 5 to 250mm. Keep the cap just above the guide hole.

[0036] Furthermore, in the above technical solution, a liquid redistribution plate with an annular shape and evenly provided through holes may be provided below the overflow weir. The liquid redistribution plate may include: an overflow ring, which is provided around the top of the through hole, and the height of the overflow ring is 1 mm to 50 mm, and is used to evenly guide the liquid phase product to the next catalyst bed; a baffle, which is provided on the side away from the distillation chamber, and the upper edge of the baffle is higher than the overflow ring.

[0037] Furthermore, in the above technical solution, the distillation chamber may include a mixing section and a separation section, and the separation section may be provided with fillers or tower plates.

[0038] Furthermore, in the above technical solution, the ratio of the diameter of the upper distillation chamber to 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 upper distillation chamber of the hydrodistillation reactor is reduced, so that the light fraction load under high pressure is fully matched with the tower plate, the tower plate separation efficiency is high, and it has full substitutability for the fractionation tower.

[0039] Furthermore, in the above technical solution, a tower plate or filler is placed in the distillation chamber, and the above filler or tower plate is a conventional form in the art, such as a filler 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, conjugated rings, flat rings, and wreaths, and the filler 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.

[0040] Furthermore, in the above technical solution, the distillation chamber plates or packing areas can be set up as reaction zones according to process requirements, in which corresponding catalysts are placed.

[0041] Furthermore, in the above technical solution, the hydrodistillation reactor further comprises: a reboiler, one end of which is connected to the outlet of the heavy oil bunker and the other end is connected to the hydrogen distribution chamber. The reboiler is used to keep the temperature of the heavy oil bunker at the temperature required by the reaction bed.

[0042] Furthermore, in the above technical solution, it is preferred to open 1 to 4 side lines in the separation section of the distillation chamber to extract the required components. At the first side line or the top of the distillation chamber, the extraction temperature 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, and the components extracted are gas and light naphtha fractions. The extraction temperature of the second side line or the side line product line is 150°C to 185°C, preferably 160°C to 170°C, and the component extracted from the side line is a heavy naphtha fraction. The extraction temperature of the third side line or the side line return line is 180 to 250°C, preferably 180 to 230°C, and the component extracted from the side line is a jet fuel fraction. The extraction temperature of the third side line or the side line return line is 200 to 370°C, preferably 210 to 350°C, and the component extracted from the side line is a diesel fraction.

[0043] Furthermore, in the hydrocracking 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 feedstock oil and the catalyst are in thin-film contact.

[0044] The hydrocracking reactor of the present invention adopts a multi-bed hydrodistillation reactor. After hydrogenation pretreatment, light aromatics in the distillate oil are separated in the separation section of the distillation chamber, wherein non-BTX monocyclic aromatic hydrocarbons with longer side chains are drawn out sideways and returned to the upper catalyst bed of the reaction chamber for cracking reaction with hydrogen, and the generated BTX-rich hydrocarbons leave the bed and enter the distiller again through the stripping chamber for separation. After the heavy distillate oil enters the cracking reaction bed, the polycyclic aromatic hydrocarbons are ring-opened to produce monocyclic aromatic hydrocarbons, which leave the reaction bed in time and are separated through the distillation chamber. The ring-opened polycyclic aromatic hydrocarbons and heavy fraction oil then enter the next bed layer. Similarly, the light aromatic fraction generated by the reaction leaves the reaction system in time and is separated out, and the heavy components of the reaction react in turn in the multiple beds below. In this way, not much of the reaction product is removed, and the cracking reaction rate is accelerated. Since the distillation chamber only separates the light aromatic products, the unseparated fractions in the distillation chamber enter the catalyst bed layer again to react. Therefore, the present invention is more selective for light aromatics and can increase the yield of light aromatics. The load of the catalyst bed is also greatly reduced. According to the test results, the load of each catalyst bed is reduced by 5% to 20%.

[0045] The catalyst bed of the multi-bed hydrocracking 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 the case in traditional reactors, will not occur.

[0046] 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.

[0047] 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.

[0048] Compared with the prior art, the present invention has the following beneficial effects: 1. The octane number or light aromatic content of 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%.

[0049] 2. The selectivity of naphtha light BTX of the present invention is higher. The present invention uses the separator or distillation chamber to produce the hydrogenation target product light naphtha immediately, 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.

[0050] 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, thus basically eliminating the coking and clogging problems of the bed.

[0051] 4. The present invention couples the reaction and separation twice. After the raw 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, which saves investment and energy consumption. At the same time, the heat generated by the hydrogenation reaction process is exactly what the distillation process needs, which also saves heat energy. The distillation process is mostly used at a relatively low pressure or vacuum. The distillation process under the conditions of medium and high pressure hydrogenation of 8MPaG to 16MPaG is generally considered to be unable 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 to achieve the expected separation effect. BRIEF DESCRIPTION OF THE DRAWINGS

[0052] Figure 1 It is a schematic diagram of the process flow of LCO diesel hydrogenation to produce more BTX according to the present invention.

[0053] Figure 2 Schematic diagrams of two structures of a liquid distributor and a gas distributor according to one embodiment of the present invention.

[0054] Figure 3 It is a schematic diagram of a catalyst bed structure according to one embodiment of the present invention.

[0055] Figure 4 It is a schematic diagram of the structure of a gas-liquid redistributor according to one embodiment of the present invention.

[0056] Figure 5 It is a schematic diagram of the structure of a gas-liquid guide plate according to one embodiment of the present invention.

[0057] Among them: 1-raw oil inlet; 2-new hydrogen; 3-hot raw oil; 4-hot hydrogen; 5-hydrogen-oil mixed raw material; 6-hydrogenation reactor; 7-bottom pipeline of hydrotreatment reactor; 8-gas-liquid separation tank; 9-multi-layer hydrodistillation reactor; 10-reaction chamber; 11-distillation chamber; 12-upper liquid distributor; 13-hydrogen; 14-upper catalyst bed; 15-lower liquid distributor; 16-side product; 17-liquid separator; 18-condenser; 19-hydrogen-rich light hydrocarbons; 20-desulfurization and deamination equipment; 21-side return line; 22-unconverted oil; 23-circulating oil pump; 24-circulating hydrogen compressor; 25-side return pump; 26-heating furnace; 27 - lower catalyst bed; 28 - stripping chamber; 101 - feed pipe; 102 - material distribution pipe; 103 - eyelet; 104, 105 - material distributor; 201 - catalyst bed; 202 - gas flow unit; 203 - gas distributor; 204 - gas distribution chamber; 205 - liquid redistribution plate; 206 - flow guide hole; 207 - flow guide groove; 208 - overflow weir; 209 - outer wall of flow guide groove; 210 - inner wall of flow guide groove; 211 - gas-liquid flow guide plate; 301, 303 - eyelet; 302 - overflow ring; 304 - baffle; 401, 501 - cap; 402 - air hole; 403 - annular cofferdam; 502 - flow guide pipe. DETAILED DESCRIPTION

[0058] 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.

[0059] 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.

[0060] 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.

[0061] 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.

[0062] like Figure 1 The schematic diagram of the process flow of hydrogenation of inferior diesel of the present invention is shown. Inferior diesel is sent into the heating furnace 24 from the raw oil inlet 1, and after being heated to 200-350°C, it merges with the heated hot hydrogen 4 to form a hydrogen-oil mixed raw material 5, and enters the hydroprocessing reactor 6. The hydroprocessing reactor is sequentially placed with a hydrorefining catalyst and a hydrocracking catalyst from top to bottom. The distillate oil after bottom hydrogenation enters the gas-liquid separation tank 8 from the bottom pipeline 7 of the hydroprocessing reactor for gas-liquid separation, and the non-condensable gas enters the condenser 18. The liquid diesel fraction enters the multi-bed hydrodistillation reactor 9, first enters the liquid distributor 15 at the bottom of the reaction chamber 10, and after passing through the liquid distribution assembly, the small molecular component hydrocarbons dissolved in the diesel are further released as light components into the distillation chamber 11, and the remaining heavy component diesel enters the catalyst bed 27 in the reaction chamber evenly. Hydrogen 13 enters from the bottom of each bed, and under the action of the gas distribution chamber 204 and the gas distributor 203, it evenly enters the catalyst bed 27. In each bed, the heavy components in the raw diesel are saturated with olefins and cracked under the action of hydrogen and catalyst to generate small molecular hydrocarbons. Under the hydrogen, the small molecular hydrocarbons and part of the large molecular hydrocarbons are separated from the bed and enter the stripping chamber 28. Part of the heavy components carried by hydrogen enter the stripping chamber 28, fall into the gas-liquid guide plate 211 under the action of gravity, and then re-enter the corresponding catalyst bed through the guide groove 207 for reaction. The light components entering the stripping chamber 28 in each bed enter the distillation chamber 11 upward, and are separated into light fractions of different temperatures under the action of the high-efficiency tower plates or fillers in the distillation chamber. The side line 16 finally obtains the target product. The side line 21 enters the catalyst bed 14 at the top of the reaction chamber through the side line return pump 25, and the light components generated after hydrogenation reaction with hydrogen enter the stripping chamber 28, and then enter the distillation chamber upward for separation again, and finally enter the side line 16 to obtain the product. The non-condensable gas at the top of the distillation chamber is further cooled by the condenser 18, and then the liquid hydrocarbon is refluxed after gas-liquid separation by the separator 17. The gas enters the desulfurization and deamination equipment 20 for purification and then is recycled by the circulating hydrogen compressor 24. The heavy diesel fraction at the bottom of the reaction chamber 10 is sent back to the inlet of the multi-bed hydrodistillation reactor 9 by the circulating oil pump 23 or extracted as a chemical raw material 22.

[0063] 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.

[0064] The experimental methods in the following examples, unless otherwise specified, are conventional methods in the art. Example 1

[0065] Using the present invention Figure 1 The flow chart shown in the figure shows that the raw oil is catalytic diesel, and its properties are shown in Table 2. After the raw oil and hydrogen are heated to 300°C in the heating furnace, they enter the hydroprocessing reactor together. The upper part of the reaction zone in the hydroprocessing reactor is catalyst A, and the lower part is filled with catalyst B. The volume ratio of the two catalysts is V A :V B =3:1, and the specific catalyst properties are shown in Table 1. The crude oil is hydrotreated in the hydroprocessing reactor, and the resulting products are separated by a gas-liquid separation tank, and the gas flows out from the top of the separation tank and flows back to the hydroprocessing reactor. The liquid is used as the raw material of the multi-bed hydrodistillation reactor. The hydrodistillation reactor is filled with catalyst C, the catalyst bed equivalent diameter-to-height ratio is 4:1, and 15 beds are set in the multi-bed hydrodistillation reactor. The opening rate of the gas guide unit below the catalyst layer is 75%, and the opening rate of the gas-liquid guide plate in the distillation chamber is 85%. The liquid raw material of the multi-bed hydrodistillation reactor is first separated from the light and heavy components in the reaction chamber through the liquid distribution component. The light component enters the distillation chamber upward and performs fine separation of different fractions according to different temperature distributions. After fine separation, the heavier fraction enters the catalyst bed downward under the action of the gas-liquid guide plate in the distillation chamber to react until the target product is generated. The heavy fractions that are not sufficiently cracked flow out from the bottom of the reaction chamber and enter the inlet of the multi-bed hydrodistillation reactor as circulating oil after passing through the circulating pump. Example 2

[0066] The difference between this embodiment and embodiment 1 is that the raw oil used is coking diesel, the catalyst bed equivalent diameter-to-height ratio in the multi-bed hydrodistillation reactor is 5:1, and the number of catalyst beds is 20. The other conditions are the same as those in embodiment 1. Example 3

[0067] The difference between this embodiment and embodiment 1 is that the raw material oil used is cycloalkyl straight-run diesel, and the catalyst bed equivalent diameter-to-height ratio is 5:1. The other conditions are the same as those in embodiment 1.

[0068] 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.

[0069] Comparative Example 1 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 separated in a separation tower and then 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 3:1, 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.

[0070] Table 1 Catalyst properties Table 2 Properties of crude oil name Catalytic diesel Coking diesel Naphthenic straight run diesel serial number 1# 2# 3# <![CDATA[Density, g / cm 3 > 0.956 0.842 0.877 Distillation range, ℃ 190～342 180～374 196～361 Sulfur content, vol% 0.95 1.2 0.13 Nitrogen content, μg / g 1300 2600 370 Table 3 Process conditions of the embodiments and comparative examples One-stage hydrogenation reactor Reaction temperature, °C 360 Reaction pressure, MPaG 8 Volume hydrogen to oil ratio 300：1 <![CDATA[Space velocity, h -1 > 1.2 Two-stage multi-bed hydrodistillation reactor Reaction temperature, °C 370 Reaction pressure, MPaG 8 Volume hydrogen to oil ratio 1200:1 <![CDATA[Space velocity, h -1 > 1.5 Table 4 Product distribution of embodiments and comparative examples

[0071] From the results in Table 4, it can be seen that the yield and selectivity of heavy naphtha are significantly higher than those of Comparative Example 1 by using the method of the present invention.

[0072] 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.

[0073] Table 5 Simulated bed temperature rise changes

[0074] It can be seen from the results in Table 5 that the temperature difference of the catalyst bed in Examples 1 to 2 of the present invention is significantly lower than that in Comparative Example 1, from 20.1°C of the traditional fixed bed to 0.4°C. The difference between the average temperature of the examples and the controlled temperature is the smallest, indicating that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. A method for hydrotreating LCO diesel to produce more BTX, comprising the following steps: (1) The raw materials LCO diesel and hydrogen enter the hydrogenation pretreatment reaction zone for hydrogenation treatment and are separated into gas and liquid; (2) The liquid produced in step (1) enters the lower bed of the hydrocracking reactor, and after being dispersed by the liquid distribution assembly, the lighter portion is carried to the middle of the reactor by the hydrogen flowing upward from the bottom, and the heavier portion enters the hydrocracking reaction zone of the multiple beds downward to contact with the hydrogen in countercurrent to carry out the hydrocracking reaction; the light products after hydrocracking in each bed of the lower hydrocracking reaction zone are separated from the reaction bed in time, and the unconverted oil enters the next bed for reaction, and the raw oil of each bed is the unconverted oil of the previous bed; the light products separated from each bed are gathered together and enter the middle of the reactor; (3) A stripping chamber is provided in the middle of the hydrocracking reaction zone, and a distillation chamber is provided in the upper part. The light components described in steps (2) and (3) first enter the stripping chamber in the middle for stripping, and then enter the distillation chamber in the upper part of the hydrocracking reactor for separation. A naphtha product rich in BTX is extracted from the side product line; (4) After the hydrogenation reaction occurs in the reactor bed in step (2), the light components produced enter the upper distillation chamber of the reactor for distillation separation, and the heavy monocyclic aromatic hydrocarbon-rich fraction extracted from the side return line enters the upper reaction zone of the hydrocracking reactor for hydrogenation reaction; the light components produced by the reaction enter the middle part of the reactor, and the heavy components enter the lower hydrocracking reaction zone; (5) The uncracked heavy components in the hydrocracking reaction zone are at least partially circulated into the hydrocracking reaction zone by a pump for reaction.

2. The hydroprocessing method according to claim 1, characterized in that: The BTX-rich naphtha fraction product in step (4) enters an aromatic extraction unit to separate light aromatic BTX products and light non-aromatic fraction oil.

3. The hydroprocessing method according to claim 1, characterized in that: The hydropretreatment reaction zone contains a hydrotreating catalyst bed; preferably, a hydrofining catalyst is placed on the upper part of the hydropretreatment reaction zone, and a hydrocracking catalyst is placed on the lower part of the hydropretreatment reaction zone, and the volume ratio of the hydrofining catalyst to the hydrocracking catalyst is 1:1 to 10:

1.

4. The hydroprocessing method according to claim 1, characterized in that: The hydrocracking reaction zone in step (2) comprises at least two catalyst beds, 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.

5. The hydroprocessing method according to claim 1 or 4, characterized in that: The catalyst bed in the hydrocracking reaction zone of step (2) is divided into at least two parts, an upper part and an lower part, each part has a liquid feed inlet, and a hydrogen feed inlet is arranged at the bottom of each bed in each part.

6. The hydroprocessing method according to claim 1, characterized in that: The operating conditions of the hydrogenation pretreatment in step (1) are as follows: the average reaction temperature is 200-450°C, the reaction pressure is 3-20 MPa, the hydrogen-to-oil volume ratio is 100:1-3000:1, and the feed oil flow rate to reactor catalyst volume ratio is 0.1-10.0 h -1 .

7. The hydroprocessing method according to claim 1, characterized in that: The average reaction temperature of the upper reaction zone of the hydrocracking reactor in step (4) is 210-430°C, the reaction pressure is 3-20 MPa, and the hydrogen-oil volume ratio is 100:1-3500:1; the average reaction temperature of the lower reaction zone of the hydrocracking reactor in step (4) is 230-430°C, the reaction pressure is 3-20 MPa, the hydrogen-oil volume ratio is 100:1-4000:1, and the ratio of the feedstock oil flow rate to the reactor catalyst volume is 0.1-10.0h -1 .

8. The hydroprocessing method according to claim 1, characterized in that: The operating conditions of the distillation chamber described in step (3) and step (4) are as follows: the reflux temperature at the top of the distillation chamber is 50°C to 70°C, and the reflux ratio is 1.1 to 10.0; the side line product extraction temperature is 150°C to 185°C, the side line return line extraction temperature is 150°C to 185°C, and the distillation chamber pressure is 3 to 20 MPa.

9. The hydroprocessing method according to claim 1, characterized in that: The diameter-to-height ratio of the catalyst bed in the reaction section of the hydrocracking reactor in step (3) is 2:1 to 20:

1.

10. A hydroprocessing device for producing more BTX from LCO diesel, the device comprising: (1) Hydrogenation pretreatment reactor and hydrotreatment reactor are used to hydrotreat the low-quality diesel raw material; (2) The hydrocracking reactor is a multi-bed hydrodistillation reactor, comprising: A reaction chamber is provided with a catalyst bed arranged in an annular manner, and the liquid raw oil and the hydrogen feed are in reverse contact in the catalyst bed; The stripping chamber is located in the middle of the hydrodistillation reactor and is a gas phase channel that runs through the reactor from top to bottom. The gas phase product after the hydrogenation reaction of each catalyst bed enters the gas phase channel under the action of stripping; The distillation chamber is arranged at the upper part of the hydrogenation distillation reactor, and the gaseous product is distilled and separated in the distillation chamber, and a return line is arranged in the middle part.

11. The hydroprocessing device according to claim 10, characterized in that: The ratio of the equivalent diameter of the cross-sectional area of ​​the catalyst bed to the total height of the catalyst bed is 0.2:1 to 10:

1.

12. The hydroprocessing device according to claim 10, characterized in that: The stripping chamber 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.

13. The hydroprocessing device according to claim 10, characterized in that: The gas-liquid guide plate is provided with a through hole, a cofferdam is provided around the through hole, and a cap is preferably provided above the through hole.

14. The hydroprocessing device according to claim 10, characterized in that: The reaction chamber is provided with 2-3 liquid feed distributors, and the liquid feed is arranged above different catalyst beds in the upper part of the reaction zone; a hydrogen feed distributor is arranged at the bottom of each catalyst bed.

15. The hydroprocessing device according to claim 14, characterized in that: The liquid phase feed distributor is of annular tube type or annular belt type, and the hydrogen feed distributor is of annular tube type or annular belt type.

16. The hydroprocessing device according to claim 10, characterized in that: A gas guide unit is provided between the hydrogen feed distributor and the bottom of the catalyst bed; the gas guide unit comprises: A guide plate is provided at the bottom of the catalyst bed and is used to receive liquid products after hydrogenation reaction from the catalyst bed; A guide pipe is evenly arranged on the guide plate and guides the hydrogen feed upward to the catalyst bed. A cap is provided on the top of the guide pipe; The overflow weir 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.

17. The hydroprocessing device according to claim 16, characterized in that: An annular liquid redistribution plate with evenly-spaced through holes is provided below the overflow weir.

18. The hydroprocessing device according to claim 16, 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 distillation chamber, the upper edge of the baffle being higher than the overflow ring.

19. The hydroprocessing device according to claim 16, characterized in that: The distillation chamber comprises a mixing section and a separation section, and the separation section is provided with fillers or tower plates.

20. The hydroprocessing device according to claim 16, characterized in that: The ratio of the diameter of the upper distillation chamber to the diameter of the lower reaction chamber is 1:1.2 to 1:

50.

21. The hydroprocessing device according to claim 16, characterized in that: The hydrodistillation reactor further comprises: a reboiler, one end of which is connected to the outlet of the heavy oil bunker, and the other end of which is connected to the hydrogen distribution chamber.

22. The hydroprocessing device according to claim 16, characterized in that: 1 to 4 side lines are opened in the separation section of the distillation chamber to extract the required components.

23. The hydroprocessing device according to claim 16, characterized in that: In the hydrocracking reactor, the diameter-to-height ratio of the catalyst bed is 3:1 to 6:1.

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