A Fischer-Tropsch synthetic oil hydroprocessing device and method

Abstract

The present invention discloses a Fischer-Tropsch synthetic oil hydroprocessing device and method. The device comprises a separator, a reaction chamber and a heavy oil bunker which are sequentially connected from top to bottom; the reaction chamber is sequentially provided with: a hydrotreating catalyst bed, a separation chamber, a hydrocracking catalyst bed and a hydrogen distribution chamber from top to bottom. The present invention fully considers the characteristics that the hydrocarbon composition, olefin and oxygen content of each fraction of the Fischer-Tropsch synthetic oil are quite different. By providing a separation chamber in the hydrotreating device, the Fischer-Tropsch synthetic oil with a wide distillation range distribution is divided into a lighter part and a heavier part. The lighter part with high olefin and oxygen content enters the hydrotreating reaction zone upward to carry out olefin saturation, deoxygenation and deacidification reaction, and the heavier part directly enters the hydrocracking reaction zone downward to carry out hydrocracking reaction, thereby avoiding the occurrence of excessive cracking of the lighter part and maximizing the yield of high-quality fuel oil.

Description

Technical Field

[0001] The present invention belongs to the field of hydrogenation reaction, and in particular, relates to a hydrogenation processing device and method for Fischer-Tropsch synthetic oil. Background Art

[0002] The Fischer-Tropsch synthesis reaction is a reaction that uses hydrogen and carbon monoxide as raw materials to produce hydrocarbons under the action of a synthesis catalyst. The Fischer-Tropsch synthesis raw material synthesis gas has a wide range of sources and can be converted from coal, natural gas, coalbed methane, biomass, etc. Fischer-Tropsch synthetic oil is mainly composed of normal alkanes, olefins and a certain amount of oxygen-containing compounds. Its non-ideal components such as sulfur, nitrogen, and aromatics are extremely low in content, and it is greatly different from conventional mineral petroleum in terms of hydrocarbon composition and main properties. The heavy naphtha fraction of Fischer-Tropsch synthetic oil basically does not contain sulfur and nitrogen, has a very low octane number (most of the olefins and alkanes are straight chains), and has a high olefin content and oxygen content; the diesel fraction has extremely low sulfur, nitrogen and aromatic content, a high cetane number, a medium olefin content and oxygen content, and a low olefin content and oxygen content in heavy oil. Therefore, Fischer-Tropsch synthetic oil must undergo corresponding hydrogenation and upgrading to obtain petrochemical raw materials or transportation fuels that meet the use specifications.

[0003] The Fischer-Tropsch synthetic oil hydrogenation and upgrading technology reported in the industry is mainly based on the technology of obtaining intermediate distillate diesel and aviation kerosene. Such as CN104711019A, CN1854265A, CN1854266A, CN101177626A, US5689031, etc. These process methods generally include two parts: hydrofining and hydrocracking. Using a fixed bed reactor, the Fischer-Tropsch synthetic oil is generally first hydrofined to saturate olefins and remove oxygen-containing compounds, and then separated into naphtha, diesel and heavy oil according to the distillation range. The heavy oil is then hydrocracking to obtain naphtha and high-quality diesel blending components. The Fischer-Tropsch synthetic oil after hydrocracking has a high content of normal alkanes in the naphtha obtained by fractionation and a high degree of saturation, which is an ideal raw material for steam cracking to produce olefins; the obtained diesel has a cetane number of 70-80, does not contain sulfur, nitrogen and aromatics, and is a very ideal component of automotive diesel. These process flows are complex and completely separate the hydrotreating and hydrocracking parts, with two fractionation systems in total, resulting in high equipment investment and high energy consumption.

[0004] Patents such as US5378348, US6309432, and CN101230291A first divide the Fischer-Tropsch synthetic oil into two or three different fractions. After the heavy fraction enters the hydroisomerization cracking reactor for isomerization cracking reaction, it is mixed with the light fraction and enters the hydrofining reactor for reaction, and the hydrogenated product enters the fractionation tower for fractionation. Although this process reduces the load of the device and simplifies the process, these inventions all use a traditional hydrogenation reaction fixed bed reaction zone, and its height-to-diameter ratio (ratio of the total height of the reactor bed to the diameter) is generally selected from 2 to 10, and the reaction temperature rise is relatively high, especially the olefin content in the Fischer-Tropsch synthetic oil is much higher than that of conventional mineral petroleum, which exacerbates the risk of "flying temperature" of the catalyst bed. In order to cool down, a large amount of cold hydrogen is injected between the beds of the reactor to reduce the reaction temperature. The reactor has high energy consumption, and in the initial stage and when the operation is abnormal, it is difficult for the cold hydrogen to take out too much heat from the reactor in time, thereby reducing the service life of the catalyst, and in severe cases, the catalyst may fail and be scrapped.

[0005] The fixed bed height-to-diameter ratio of traditional hydrogenation reaction is relatively high to ensure sufficient contact between gas-liquid materials and solid catalysts to achieve the required reaction depth and efficiency. Dong Fangliang et al. mentioned in "Determination of the Main Structural Parameters of Hydrogenation Reactors" 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 traditional fixed bed reactors is mostly 4 to 9". Patent CN109679689 A also mentioned that the height-to-diameter ratio of existing liquid phase hydrogenation reactors 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 confirmed that the design is reasonable and has a relatively general adaptability. The extensive industrial success may also lead to the failure of technicians 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 there are only research reports that prove that the existing height-to-diameter ratio is a suitable design.

[0006] In addition, Fischer-Tropsch synthetic oil is mainly composed of straight-chain high-carbon number alkanes and olefins, especially the light fraction has a higher content of heavy olefins, which can reach more than 50%, while traditional petroleum-based catalytic cracking raw materials are mainly composed of polycyclic hydrocarbons with side chains, and the molecular structures of the two are very different. Compared with polycyclic hydrocarbon molecules with side chains, long-chain hydrocarbon molecules have lower cracking activation energy and are more likely to undergo cracking reactions. Therefore, during the cracking process, cracking products may undergo cracking reactions again, making the reaction depth uncontrollable, producing a large amount of gas and olefins, and causing coking and blockage in the reactor. However, none of the existing methods can eliminate the coking and blockage phenomenon of traditional reactors. This is because the raw oil and hydrogen products cannot leave the system during the flow of the entire reactor bed, and the materials react and flow downward during the bed advancement process. During this period, both fresh materials react under the action of the catalyst and cracking products undergo cracking reactions again. Summary of the invention

[0007] In view of the shortcomings of the prior art, the purpose of the present invention is to provide a Fischer-Tropsch oil hydroprocessing device and method, so as to effectively control the reaction depth, greatly improve the yield of heavy naphtha and diesel, eliminate the risk of "overheating" of the bed, and extend the service life of the catalyst.

[0008] The first aspect of the present invention is to provide a Fischer-Tropsch oil hydroprocessing device, the device comprising:

[0009] (1) A separator, a reaction chamber and a heavy oil tank connected in sequence from top to bottom;

[0010] (2) The reaction chamber is provided with: a hydrorefining catalyst bed, a separation chamber, a hydrocracking catalyst bed and a hydrogen distribution chamber from top to bottom.

[0011] In the above technical solution, the ratio of the equivalent diameter of the reaction chamber cross-sectional area (the equivalent diameter formula is de=4A / L, A is the cross-sectional area of ​​the reaction chamber, and L is the circumference of the bed) to the height is 2:1 to 10:1, preferably 3:1 to 6:1.

[0012] In the above technical solution, the height of the reaction chamber is generally 100 mm to 5000 mm, preferably 200 mm to 1000 mm. At a higher diameter-to-height ratio, the material flux through the catalyst bed in the reaction chamber can be greatly increased, while reducing the residence time of the material and heat in the catalyst bed.

[0013] In the above technical solution, the cavity of the reaction chamber can be a horizontal tank or a cylindrical form, preferably a horizontal tank form. There are heads on both sides of the cavity to facilitate loading and unloading of the catalyst. The reaction chamber is filled with catalyst to form a reaction zone, and the reaction zone is divided into multiple reaction zone units by a mesh partition.

[0014] In the above technical solution, the height of the separation chamber arranged in the reaction chamber is generally 10% to 30% of the height of the reaction chamber. The ratio of the height of the hydrorefining catalyst bed to the height of the hydrocracking catalyst bed is 5:1 to 1:8, preferably 2:1 to 1:6.

[0015] In the above technical solution, a liquid distribution assembly is arranged in the separation chamber, and the liquid distribution assembly 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 hydrogenation 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 naphtha and diesel fractions, and the heavier part is usually heavy diesel and wax oil fractions.

[0016] In the above technical scheme, the liquid distributor is a conventional distributor in the art, such as a showerhead distributor, a coil distributor, a porous straight tube distributor, a straight tube baffle distributor, a baffle distributor, a tangential circulation distributor, a rotating blade distributor, a double-row blade distributor, etc. In the present invention, the liquid distributor is preferably a porous tube distributor and a straight tube baffle distributor, and the pore diameter of the tube distributor is 0.5 to 20 mm, preferably 2 to 10 mm. The farther from the inlet end of the crude oil, the larger the pore size. The height of the liquid distributor from the top of the reactor bed is 1 to 1000 mm, preferably 50 to 500 mm. The height is related to the properties, temperature and pressure of the crude oil. Generally speaking, when the temperature is higher, the liquid distributor is higher from the bed, so that the distributor can make the raw materials fall more evenly on the bed surface in a higher space. Similarly, the higher the pressure, the larger the injection angle of the liquid distributor, the lower the height from the top of the reactor bed, and the more space is saved.

[0017] In the above technical solution, the shape of the liquid distribution plate is the same as the cross-section of the catalyst bed in the hydrogenation main reactor, and the area of ​​the liquid distribution plate is 10% to 100% of the cross-section area of ​​the catalyst bed, preferably 60% to 100%.

[0018] In the above technical solution, a plurality of first through holes are evenly arranged on the liquid 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 first through hole is 5 to 100 mm, and the height of the first overflow ring is 1 to 30 mm.

[0019] In the above technical solution, 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.

[0020] In the above technical solution, the distribution cone is arranged at the upper center of the liquid distribution plate, the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through 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 second 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.

[0021] In the above technical solution, the ratio of the separator diameter or equivalent diameter to the diameter or equivalent diameter of the lower reaction chamber is 1:1.2 to 1:50, preferably 1:2 to 1:10. The upper separator diameter is reduced, so that the light fraction load under high pressure is completely matched with the tower plate, the tower plate separation efficiency is high, and it has complete substitutability for the fractionation tower.

[0022] In the above technical solution, the separator is composed of a mixing section, a separation section and a stabilization section from bottom to top. The height of the mixing section is 20% to 35% of the total height of the separator, the height of the separation section is 55% to 70% of the total height of the separator; the height of the stabilization section is 5% to 10% of the total height of the separator.

[0023] In the above technical scheme, packing or tower plates are placed in the separation section. The packing or tower plates are all conventional forms in the art, such as the packing can be selected from one or more of the 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 packing can also be selected from metal or ceramic corrugated packing. The tower plate is one or more of the tower plates with downcomers such as bubble plates, sieve plates, float valve plates, mesh plates, tongue-shaped plates, guide sieve plates, and multi-downcomer tower plates, and can also be tower plates without downcomers such as flow-through sieve plates and flow-through corrugated plates. High-efficiency tower plates such as guide float valves and sieve plate towers are preferred. There is no restriction on whether to place packing in the mixing section and the stabilization section, and the reaction zone can be increased according to the process requirements.

[0024] In the above technical solution, multiple partitions are arranged in parallel in the vertical direction in the hydrogen distribution chamber and the catalyst bed in the reaction chamber, and the multiple partitions divide the hydrogen distribution chamber into multiple air intake units, and at least one hydrogen inlet is arranged at the bottom of each air intake unit. Multiple 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%.

[0025] In the above technical solution, a gas distributor is provided at the hydrogen inlet. In the present invention, the gas distributor is preferably a tangential circulation distributor or a rotating blade distributor, which can make the flow rate of the gas entering the entire catalyst bed interface more uniform, avoiding bias flow, channel flow and the like.

[0026] The hydrogen distribution chamber of the hydroprocessing device of the present invention is provided with a hydrogen feed pipe, and the hydrogen feed pipe is provided with multiple inlets, each hydrogen feed pipe corresponds to the catalyst bed area between the two partitions, so that after hydrogen comes out of each distributor, it can pass upward through the top reaction zone. There is at least one channel at the connection between the perforated baffle and the bottom of the main reactor.

[0027] In the above technical solution, there are 1 to 3 gas distributors at the bottom of the catalyst bed between each two adjacent partitions, and the distribution area of ​​hydrogen from all gas distributors in the partition area when it reaches the bottom of the bed in the corresponding area should cover the entire area. Further, the partition is annular or circular.

[0028] In the above technical solution, the heavy oil bin is arranged at the bottom center of the reaction chamber, and the heavy oil bin is connected to the multiple air inlet units.

[0029] In the above technical solution, the hydroprocessing device 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.

[0030] In the above technical solution, the hydroprocessing device also includes: a multi-stage auxiliary reaction chamber, each stage of the auxiliary reaction chamber is independently fed with hydrogen, a heavy oil tank is independently arranged at the bottom center, the liquid raw material inlet of each stage of the auxiliary reaction chamber is connected to the heavy oil tank of the previous stage, and the tops of the multi-stage auxiliary reaction chambers are all connected to the separator.

[0031] The second aspect of the present invention is to provide a Fischer-Tropsch synthetic oil hydroprocessing method, comprising the following steps:

[0032] (1) The Fischer-Tropsch synthetic oil enters the separation chamber in the hydroprocessing unit, where the lighter part is carried away by the hydrogen flowing upward from the bottom and enters the hydrofining reaction zone upward; the heavier part enters the hydrocracking reaction zone downward and undergoes a cracking reaction with the countercurrent hydrogen. The light components produced by the reaction separate upward from the hydrocracking reaction zone and enter the hydrofining reaction zone, while the heavy components flow downward for reflux or are partially thrown out and discharged as tail oil;

[0033] (2) The hydrofinished material flows upward into a separator disposed at the top of the hydrofinishing reaction zone, where it is separated into naphtha and diesel fractions. The separated uncracked heavy components flow downward into the hydrocracking reaction zone again for hydrocracking reaction.

[0034] In the above technical solution, after the Fischer-Tropsch synthetic oil enters the separation chamber in step (1), it is preferably dispersed into a lighter part and a heavier part through flash evaporation and a liquid distribution component.

[0035] In the above technical scheme, the ratio of the equivalent diameter of the catalyst bed in the hydrotreating reaction zone and the hydrocracking reaction zone in step (1) (the equivalent diameter formula is de=4A / L, A is the cross-sectional area of ​​the bed, and L is the circumference of the bed) to the total height of the catalyst bed is 2:1 to 10:1, preferably 3:1 to 6:1.

[0036] 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 heights of multiple catalyst beds. The bed cross-sectional area described in the present invention refers to the cross-sectional area of ​​the reactor bed. The reactor bed is preferably isodiametric, 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-down 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 at various locations or the cross-sectional area of ​​the reaction chamber within the catalyst bed range.

[0037] The Fischer-Tropsch synthesis oil of the present invention includes high-temperature Fischer-Tropsch synthesis full-fraction oil and low-temperature Fischer-Tropsch synthesis full-fraction oil, preferably low-temperature Fischer-Tropsch synthesis full-fraction oil. The properties of the Fischer-Tropsch synthesis oil are as follows: density 0.6 g / cm 3 ~1.0g / cm 3 , preferably 0.7 g / cm 3 ~0.95g / cm 3 ; Final distillation point 650℃~750℃, preferably 680℃~720℃.

[0038] In the above technical solution, the hydrotreating reaction zone described in step (1) is filled with a hydrotreating catalyst. The hydrotreating catalyst is generally based on alumina or silicon-containing alumina as a carrier and a metal of group VIB and group VIII as an active component, such as one, two or more selected from W, Mo, Co, and Ni.

[0039] In the above technical scheme, the hydrocracking reaction zone described in step (1) is loaded with a hydrocracking catalyst. The catalyst generally comprises an active component and a carrier, the carrier component comprises one or more of alumina, silicon-containing alumina and molecular sieve, preferably a molecular sieve, and the molecular sieve may be a Y-type molecular sieve agent; the active component is one or more of Group VIB and Group VIII metals, the Group VIB metals are generally Mo and / or W, and the Group VIII metals are generally Co and / or Ni.

[0040] The shapes of the hydrorefining catalyst and the hydrocracking catalyst can be any conventional existing hydrocracking catalyst shapes, preferably porous catalysts, shaped catalysts and / or honeycomb catalysts. The porous catalyst pore diameter 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 honeycomb catalyst pore diameter or pore side length is 1 to 50 mm, preferably 3 to 15 mm; the catalyst bed void ratio is recommended to be 15% to 85%, preferably 20% to 75%.

[0041] In the above technical scheme, the operating conditions of the Fischer-Tropsch synthetic oil hydrogenation reaction process (including hydrofining and hydrocracking) described in step (1) are as follows: reaction temperature is 380°C to 450°C, reaction pressure is 3MPa to 15MPa, top reflux ratio is 1.2 to 4.5, hydrogen-to-oil volume ratio is 300 to 2000, liquid hourly volume space velocity is 0.1h -1 ~10.0h -1 The preferred operating conditions are: reaction temperature of 400°C to 450°C, reaction pressure of 4MPa to 12MPa, top reflux ratio of 1.5 to 3.0, hydrogen to oil volume ratio of 400 to 1500, liquid hourly volume space velocity of 0.5h -1 ~10.0h -1 .

[0042] In the above technical solution, the separator extracts products at the side line or the top. Gas is extracted at the top, and naphtha and diesel fractions are extracted at the side line.

[0043] In the above technical solution, it is preferred to open 2 to 3 side lines in the separation section of the separator to extract the required products. At the top of the separator, the extraction temperature is 60°C to 80°C, and the components extracted are gas and light naphtha fraction. The extraction temperature of the first side line is 150°C to 190°C, preferably 160°C to 180°C, and the component extracted from this side line is the heavy naphtha fraction. The extraction temperature of the second side line is 300°C to 380°C, preferably 310°C to 370°C, and the component extracted from this side line is the diesel fraction.

[0044] Furthermore, in the above technical solution, the side line extraction line can be provided with a reflux.

[0045] After a lot of research, it was found that for the gas-liquid-solid three-phase reaction process in which the amount of liquid phase decreases rapidly and the amount of gas phase increases rapidly during the reaction, the rapid increase in the amount of gas phase occupies a large amount of bed space, which greatly increases the liquid phase flow rate. According to the traditional design, although the gas-liquid-solid three-phase contact can be guaranteed, the effective reaction time of the liquid phase that needs further conversion is reduced, and the probability of the gas phase that does not need to react again (such as the gas phase obtained by liquid phase conversion under reaction conditions) contacting the catalyst increases. For the system that requires more liquid phase conversion and gas phase control of secondary reaction, the overall reaction effect is subject to certain limitations, which is generally manifested as the reaction conversion rate, selectivity, etc. are difficult to further improve.

[0046] 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 target product heavy naphtha in the inferior diesel hydrogenation technology), while solving the problems of easy flooding of the countercurrent reactor and limited hydrogen-to-oil ratio.

[0047] Compared with the prior art, the advantages of the present invention are:

[0048] (1) The present invention fully considers the characteristics of the large differences in hydrocarbon composition, olefin and oxygen content of each fraction of Fischer-Tropsch synthetic oil. By setting a separation chamber in the hydroprocessing device, the Fischer-Tropsch synthetic oil with a wide distillation range distribution is divided into a lighter part and a heavier part. The lighter part with high olefin and oxygen content enters the hydrorefining reaction zone upward for olefin saturation, deoxygenation and deacidification reaction, and the heavier part directly enters the hydrocracking reaction zone downward for hydrocracking reaction, thereby reducing the amount of hydrorefining and hydrocracking catalyst, avoiding the occurrence of excessive cracking of the lighter part on the hydrocracking catalyst, and maximizing the yield of high-quality fuel oil. At the same time, the light product produced after the hydrocracking of the heavier part contains a small amount of olefins, and the olefins therein can be hydrogenated and saturated by the hydrorefining catalyst upward, thereby ensuring the stability of product quality.

[0049] (2) The design of the liquid distribution assembly can disperse the heavy component liquid flowing downward in the separation chamber into suitable small droplets, 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 assembly allows the heavy component 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 reactor of the present invention, under the same process conditions and product index requirements, the porosity of the reactor of the present invention can be smaller. 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 and ensure good mass transfer performance. Finally, the product properties have good controllability.

[0050] (3) The present invention can realize the timely extraction of naphtha and diesel products through the arrangement of a separator, flash evaporation and stripping, which not only takes away a large amount of heat released by the reaction, but also controls the degree of reaction, prevents excessive cracking and gasification of naphtha and diesel products, and guarantees the yield of naphtha and diesel products to the maximum extent. At the same time, since the product partial pressure is always kept at a low state, the reaction speed is accelerated, the hidden danger of by-products clogging the catalyst is eliminated, the yield of the target product is improved, and the service life of the catalyst is extended. BRIEF DESCRIPTION OF THE DRAWINGS

[0051] Figure 1 1 is a schematic diagram of the structure of the Fischer-Tropsch oil hydroprocessing device of the present invention;

[0052] Figure 2 It is a schematic diagram of the process flow of the hydroprocessing of Fischer-Tropsch synthetic oil and a schematic diagram of the structure of the hydroprocessing device of the present invention;

[0053] Among them: 1-Fischer-Tropsch synthesis oil; 2-hydrogen; 3-hot raw oil; 4-hot hydrogen; 5-hydrogenation reactor inlet pipeline; 6-hydrogenation reactor reaction chamber; 7-hydrogenation reactor heavy oil bunker; 8-hydrogenation reactor separation chamber; 9-mixing section isomerization decondensation catalyst bed; 10-stabilization section; 11-hydrogenation refining catalyst bed; 12-hydrogenation cracking catalyst bed; 13-separation section; 14-liquid distribution component; 15-mixing section; 16-separator; 17-separator tower top condenser; 18-liquid separation tank; 19-hydrogen-rich and non-condensable gas; 20-heavy naphtha product; 21-diesel product; 22-light naphtha fraction; 23-hydrogenation heavy fraction; 24-reflux heavy fraction; 25-circulating oil pump; 26-tail oil; 27-mesh baffle; 28-heating furnace.

[0054] Figure 3 It is a schematic side view of the structure of a circular segment according to one embodiment of the present invention.

[0055] Figure 4 It is a schematic side view of the structure of a notch partition according to another embodiment of the present invention. The partition is only located in the hydrogen distribution chamber, and there is no notch partition inside the catalyst bed.

[0056] Figure 5 1 is a schematic diagram of a top view of the structure of an annular partition according to an embodiment of the present invention.

[0057] FIG6 is a schematic side view of the annular partition according to two embodiments of the present invention. Figure 6-2 The partition is only located in the hydrogen distribution chamber, and there is no annular partition inside the catalyst bed.

[0058] Figure 7 It is a schematic diagram of the top structure of a liquid distribution assembly according to one embodiment of the present invention.

[0059] Figure 8 It is a schematic diagram of the top view and side view of the structure of a liquid distribution assembly according to another embodiment of the present invention.

[0060] Fig. 9 3 is a schematic diagram of a top view of the structure of a first overflow ring according to an embodiment of the present invention.

[0061] Fig.10 yes Figure 8 Schematic diagram of the three-dimensional structure of the first overflow ring. DETAILED DESCRIPTION

[0062] The present invention will be further described below in conjunction with the accompanying drawings.

[0063] like Figure 1 and Figure 2 The Fischer-Tropsch synthetic oil hydroprocessing device and process flow of the present invention are shown. The Fischer-Tropsch synthetic oil 1 is fed into the heating furnace 28, heated to 380°C to 450°C to obtain hot raw oil 3, which enters the hydrogenation reactor separation chamber 8 in the hydrogenation reactor reaction chamber 6 through the hydrogenation reactor inlet pipeline 5, and after flash evaporation and dispersion by the liquid distribution component 14, the lighter part is carried away by the hydrogen flowing upward from the bottom and enters the hydrorefining catalyst bed 11 (including 11-1 to 11-6); the heavier part is sprayed downward evenly on the hydrocracking catalyst bed 12 (including 12-1 to 12-6). At the same time, the hot hydrogen 4 heated by the heating furnace 28 is sent into the reaction chamber 6 of the hydrogenation reactor, and under the action of the mesh baffle 27 and the gas distributors 4-1 to 4-6, it moves upward from the bottom of each reaction zone evenly, and contacts with the heavier part sprayed from the liquid distribution component in the hydrocracking catalyst bed 12 in countercurrent, and flows upward with the lighter part flowing upward from the separation chamber in the hydrorefining catalyst bed 11. Under the operating pressure of 4MPa to 15MPa, the heavier components of the Fischer-Tropsch synthetic oil undergo hydrocracking reaction on the hydrocracking catalyst bed, and some long-chain molecules are broken into short-chain molecules, and some polycyclic aromatic hydrocarbons are also broken; the lighter components of the Fischer-Tropsch synthetic oil, the light components of the cracking reaction and the hydrogen undergo olefin hydrogenation saturation, hydrodeoxygenation and impurity removal reactions on the hydrorefining catalyst bed 11.

[0064] Smaller molecular hydrocarbons or other gases produced by the hydrogenation reaction are quickly brought to the separator 16 by hydrogen. After the separation action of the mixing section 15 and the separation section 13, a part of the heavier fractions pass downward through the hydrotreating catalyst bed 11, the hydrogenation reactor separation chamber 8 and the liquid distribution assembly 14 and evenly fall onto the surface of the hydrocracking catalyst bed 12. Another part of the light components continue to move to the top stabilizing section 10 in the separator, flow out from the top of the separator 16, continue to be cooled by the separator tower top condenser 17, and then pass through the liquid separator 18 for gas-liquid separation. The hydrogen-rich and non-condensable gas 19 are discharged, and the liquid hydrocarbons are refluxed or extracted as the light naphtha fraction 22. The heavy naphtha product 20 is extracted at the position of 150°C to 190°C of the separator side line, and the diesel product 21 is extracted at the position of 300°C to 380°C of the separator side line.

[0065] The hydrogenated heavy fraction 23 in the hydrogenation reactor flows out from the heavy oil tank 7 of the hydrogenation reactor, and is mixed with the raw oil as the feed of the hydrogenation reactor through the circulating oil pump 25 as the reflux heavy fraction 24, and can also be partially discharged from the device as tail oil through the circulating oil pump 25.

[0066] Further, in one or more exemplary embodiments of the present invention, the hydrogenation reactor reaction chamber 6 may be a horizontal storage tank, such as Figure 1 As shown, its axial direction is arranged in the horizontal direction, and the ends of the horizontal storage tank are provided with heads. Further, in one or more exemplary embodiments of the present invention, the reaction chamber 6 can also be a flat cylindrical tank, and its axial direction is arranged in the longitudinal direction.

[0067] Further, in one or more exemplary embodiments of the present invention, the shape of the partition matches the bottom of the reaction chamber 6. When the reaction chamber 6 is a horizontal storage tank, the partition is a circular partition, such as Figure 3 and Figure 4 When the reaction chamber 6 is a flat cylindrical tank, the plurality of partitions are coaxial annular partitions, such as Figure 5 to Figure 7 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 1 As shown, a plurality of mesh baffles 27 can extend upward to the hydrofining catalyst bed 11, and the opening rate of the baffles at the bottom that are not in contact with the catalyst bed 11 is less than 70%. The lower opening rate is conducive to increasing resistance, so that hydrogen can enter the hydrofining catalyst bed 11 upward as much as possible, further playing the role of a gas distributor. The opening rate of the baffles in the hydrofining catalyst bed 11 is greater than 50%, which is conducive to more effective use of the catalyst.

[0068] Further, in one or more exemplary embodiments of the present invention, the separator 16 includes a mixing section 15 , a separation section 13 , and a stabilization section 10 from bottom to top.

[0069] Further, in one or more exemplary embodiments of the present invention, the liquid distribution assembly 14 disposed at the bottom of the separation chamber 8 of the hydrogenation reactor includes a liquid distribution plate and a distribution cone. The liquid distribution plate is disposed 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 first through holes are evenly provided on the liquid distribution plate, and a first overflow ring is provided around the first through hole, and an overflow portion is provided at the outer edge of the distribution plate (not shown in the figure). The distribution cone is disposed at the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring is provided around the second through hole (not shown in the figure).

[0070] Preferably, but not limiting, in one or more exemplary embodiments of the present invention, in combination with Fig. 9 and Fig.10 As shown, a sawtooth portion 14-1 is provided on the inner side of the first overflow ring, the sawtooth portion is bent downward, and a guide groove 14-2 is provided on the sawtooth portion. Exemplarily, the guide groove is opened along the center of the sawtooth portion.

[0071] In one or more embodiments of the present invention, the hydrogenation reactor further comprises an auxiliary reaction chamber. It should be understood that the auxiliary reaction chamber can be multi-stage. Each stage of the auxiliary reaction chamber is independently fed with hydrogen, and a heavy oil bunker is independently arranged at the bottom center. The liquid raw material inlet of each stage of the auxiliary reaction chamber is connected to the heavy oil bunker of the previous stage, and the tops of the multi-stage auxiliary reaction chambers are all connected to the separator 16.

[0072] The present invention is further described below in conjunction with specific embodiments, but it should be understood that the protection scope of the present invention is not limited by the specific implementation methods.

[0073] Example 1

[0074] Using the present invention Figure 1 and Figure 2 The device and process flow shown in the figure, the raw oil is Fischer-Tropsch synthetic oil, and the properties are shown in Table 1. After the raw oil and hydrogen are heated to 430°C in the heating furnace, they enter the hydrogenation reactor, in which the hydrorefining catalyst FHUDS-5 bed and the hydrocracking catalyst FC-14 bed produced by Sinopec Dalian Petrochemical Research Institute are placed from top to bottom. The equivalent diameter-to-height ratio of each catalyst bed is 4:1, and the height of the hydrorefining catalyst bed and the height of the hydrocracking catalyst bed are 800 mm respectively. The height of the separation chamber between the hydrorefining catalyst bed and the hydrocracking catalyst bed is 15% of the height of the reaction chamber. An annular partition is arranged in the catalyst bed, and the number of partitions is 4. 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%.

[0075] The raw liquid first passes through the liquid distribution assembly to separate the light and heavy components. The light components are then hydrofined and then enter the separator. The heavy components enter the four different sub-catalyst bed areas in the hydrocracking catalyst bed, and react with the hydrogen entering from the bottom under the action of the hydrocracking catalyst to undergo hydrocracking, olefin saturation and other reactions. The light fractions produced by the reaction quickly separate from the reaction system upwards, and then enter the separator after hydrofining. The separator is divided into a mixing section, a separation section and a stabilization section from bottom to top. The height of the mixing section is 30% of the total height of the separator, the height of the separation section is 65% of the total height of the separator, and the height of the stabilization section is 5% of the total height of the separator. Raschig rings are placed in the mixing section and the stabilization section. After the hydrofining logistics is mixed with the light components in the lower part of the mixing section of the separator, it continues to move upward. Under the action of the separation tower plates in the mixing section and the separation section, the heavy fraction re-enters the hydrocracking catalyst bed for cracking reaction. The light fraction is extracted from the side line as heavy naphtha fraction and diesel fraction and sent out as products. The non-condensable gas continues to be cooled through the condenser, and after gas-liquid separation in the liquid separator, liquid hydrocarbons are extracted as light naphtha fractions. The gas enters the desulfurization and deamination equipment for purification and recycling. The heavy fraction that is not cracked enough flows out from the bottom of the heavy oil tank and passes through the circulation pump and then enters the inlet of the hydroprocessing unit as circulating oil. The specific operating process conditions are shown in Table 2, and the product distribution and properties are shown in Table 4.

[0076] The liquid distribution assembly includes a liquid distributor, a liquid distribution plate and a distribution cone arranged above the liquid distributor. The liquid distribution plate has the same shape as the top surface of the catalyst bed, and the area of ​​the liquid distribution plate is 70% of the cross-section of the catalyst bed. The liquid distribution plate is evenly provided with a plurality of first through holes, a first overflow ring is arranged around the first through hole, an overflow portion is arranged at the outer edge of the distribution plate, the opening rate of the liquid distribution plate is 50%, the diameter of the first through hole is 10 mm, and the height of the first overflow ring is 10 mm. The distribution cone is arranged at the upper center of the liquid distribution plate, the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through hole; the top angle of the distribution cone is 120°, the opening rate of the distribution cone is 50%, and the height of the second overflow ring is 10 mm; the bottom area of ​​the distribution cone is 10% of the area of ​​the liquid distribution plate.

[0077] Example 2

[0078] The difference between this embodiment and embodiment 1 is that the equivalent diameter-to-height ratio of the hydrorefining and hydrocracking catalyst beds is 5:1, respectively, and the number of partitions in the catalyst bed is 6. 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 30%, and the opening rate of the partition in the catalyst layer is 80%. The area of ​​the liquid distribution plate is 90% of the cross-section of the catalyst bed. The opening rate of the liquid distribution plate is 80%, the diameter of the first through hole is 20mm, and the height of the first overflow ring is 20mm. The distribution cone is arranged at the upper center of the liquid distribution plate, and the distribution cone is provided with multiple second through holes, and a second overflow ring is arranged around the second through hole; the top angle of the distribution cone is 150°, the opening rate of the distribution cone is 70%, and the height of the second overflow ring is 20mm; the bottom area of ​​the distribution cone is 15% of the area of ​​the liquid distribution plate. The remaining conditions are the same as those in embodiment 1.

[0079] Example 3

[0080] The difference between this embodiment and embodiment 1 is that the equivalent diameter-to-height ratio of the hydrorefining and hydrocracking catalyst bed is 6:1, and the number of partitions in the catalyst bed is 6. The area of ​​the liquid distribution plate is 80% of the cross section of the catalyst bed. The opening rate of the liquid distribution plate is 70%, the diameter of the first through hole is 30 mm, and the height of the first overflow ring is 20 mm. The distribution cone is arranged at the upper center of the liquid distribution plate, and the distribution cone is provided with a plurality of second through holes, and a second overflow ring is arranged around the second through hole; the top angle of the distribution cone is 140°, the opening rate of the distribution cone is 60%, and the height of the second overflow ring is 20 mm; the bottom area of ​​the distribution cone is 15% of the area of ​​the liquid distribution plate. The remaining conditions are the same as those in embodiment 1.

[0081] Example 4

[0082] The difference between this embodiment and embodiment 1 is that the height of the mixing section of the separator is 25% of the total height of the separator, the height of the separation section is 65% of the total height of the separator, and the height of the stabilization section is 10% of the total height of the separator. The other conditions are the same as those of embodiment 1.

[0083] Comparative Example 1

[0084] A conventional two-stage hydrogenation process, i.e., a refining + cracking process, is adopted. Both the refining and cracking reactors adopt a reaction process in which the raw material and hydrogen flow in parallel from top to bottom. The Fischer-Tropsch synthetic oil is first subjected to olefin saturation, deoxygenation and impurity removal reactions in the hydrorefining reactor, and the hydrorefining logistics enters the first separation and fractionation system to obtain refined high-pressure gas, naphtha, diesel and heavy oil, wherein the refined high-pressure gas refluxes into the hydrorefining reactor, and the heavy oil enters the hydrocracking reactor for hydrocracking reaction. After the reaction is completed, it enters the second separation and fractionation system, and gas, naphtha and diesel fractions are separated. The uncracked heavy components after separation are all refluxed into the hydrocracking reactor for hydrocracking reaction again. The catalyst used is the same as that in Example 1, and the process conditions are shown in Table 3.

[0085] Table 1 Main properties of crude oil

[0086] project data <![CDATA[Density, g / cm 3 > 0.815 Distillation range, ℃ Initial distillation point / 10% 65 / 203 30％ / 50％ 312 / 371 70％ / 90％ 460 / 557 95％ 617 Sulfur content, vol% 4.1 Nitrogen content, μg / g 4.9 Oxygen content, wt% 0.82

[0087] Table 2 Example process conditions

[0088] Example 1 Reaction temperature, °C 410 Reaction pressure, MPa 8.0 Volume hydrogen to oil ratio 1000：1 <![CDATA[Space velocity, h -1 > 1.5

[0089] Table 3 Process conditions of Comparative Example 1

[0090] Example 1 One-stage hydrotreating reactor Reaction temperature, °C 300 Reaction pressure, MPa 8.0 Volume hydrogen to oil ratio 300：1 <![CDATA[Space velocity, h -1 > 3.0 Secondary hydrocracking Reaction temperature, °C 410 Reaction pressure, MPa 8.0 Volume hydrogen to oil ratio 1000:1 <![CDATA[Space velocity, h -1 > 1.5

[0091] Table 4 Product distribution and properties

[0092] Example 1 Example 2 Example 3 Example 4 Comparative Example 1 Naphtha fraction Distillation range, ℃ 65～160 65～160 65～160 65～160 65～160 Yield, wt% 32.2 31.3 33.8 32.3 18.4 <![CDATA[Density, g / cm 3 > 0.708 0.705 0.710 0.706 0.714 Composition, wt% Normal alkanes 91.20 92.30 91.57 91.74 87.57 Isoalkanes 6.54 6.52 6.81 7.02 10.24 Diesel fraction Distillation range, ℃ 160～370 160～370 160～370 160～370 160～370 Yield, wt% 66.1 67.2 65.0 66.8 78.1 <![CDATA[Density, g / cm 3 > 0.781 0.773 0.765 0.780 0.793 Pour point, ℃ -3 -4 -5 -3 -2 Cetane number 83 83 84 83 78

[0093] From the results in Table 4, it can be seen that the yield of diesel by the method of the present invention is significantly higher than that of Comparative Example 1, and the cetane number is greater than 83, which is a high-quality diesel blending component; the naphtha fraction is mainly composed of alkanes and is a high-quality raw material for steam cracking to produce ethylene and propylene.

[0094] Example 5

[0095] 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.

[0096] Table 5 Simulated bed temperature rise changes

[0097] Bed temperature point Example 1 Example 2 Comparative Example 1 Maximum radial temperature difference, ℃ 3.1 3.5 28.1 Average temperature, °C 411.2 411.5 431.6

[0098] 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 28.1°C of the traditional fixed bed to 3.5°C. The difference between the average temperature of the examples and the controlled temperature is small, indicating that the reactor of the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.

Claims

1. A Fischer-Tropsch oil hydrotreating device, It is characterized in that The device comprises: (1) A separator, a reaction chamber and a heavy oil tank connected in sequence from top to bottom; (2) The reaction chamber is provided with: a hydrorefining catalyst bed, a separation chamber, a hydrocracking catalyst bed and a hydrogen distribution chamber from top to bottom; The ratio of the equivalent diameter of the cross-sectional area of ​​the reaction chamber to the height is 2:1 to 10:1; A plurality of partitions are arranged in parallel in the vertical direction in the hydrogen distribution chamber and the hydrorefining 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; and a plurality of holes are distributed on each partition; The separator consists of a mixing section, a separation section and a stabilization section from bottom to top; The heavy oil bin is arranged at the bottom center of the reaction chamber, and the heavy oil bin is connected with the plurality of air inlet units; A liquid distribution assembly is arranged in the separation chamber, and the liquid distribution assembly comprises a liquid distributor and a liquid distribution plate and a distribution cone arranged above the liquid distributor; The area of ​​the liquid distribution plate is 10% to 100% of the cross-sectional area of ​​the catalyst bed; the cross-sectional area of ​​the catalyst bed is the cross-sectional area of ​​the reaction chamber; The opening rate of the liquid distribution plate is 5% to 90%, the diameter of the first through hole is 5 to 100 mm, and the height of the first overflow ring is 1 to 30 mm; The top angle of the distribution cone is greater than 90°, the opening rate of the distribution cone is 5% to 80%, and the height of the second 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; 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; The distribution cone is arranged at the upper center of the liquid distribution plate, and a plurality of second through holes are arranged on the distribution cone, and a second overflow ring is arranged around the second through holes.

2. The device according to claim 1, It is characterized in that The ratio of the equivalent diameter of the cross-sectional area of ​​the reaction chamber to the height is 3:1 to 6:

1.

3. The device according to claim 1, It is characterized in that The height of the reaction chamber is 100 mm to 5000 mm.

4. The device according to claim 3, It is characterized in that The reaction chamber is 200 mm to 1000 mm high.

5. The device according to claim 1, It is characterized in that The height of the separation chamber arranged in the reaction chamber is 10% to 30% of the height of the reaction chamber; the ratio of the height of the hydrorefining catalyst bed to the height of the hydrocracking catalyst bed is 5:1 to 1:

8.

6. The device according to claim 5, It is characterized in that The ratio of the hydrorefining catalyst bed height to the hydrocracking catalyst bed height is 2:1 to 1:

6.

7. The device according to claim 1, It is characterized in that The area of ​​the liquid distribution plate is 60% to 100% of the cross-sectional area of ​​the catalyst bed.

8. The device according to claim 1, It is characterized in that The ratio of the separator diameter or equivalent diameter to the lower reaction chamber diameter or equivalent diameter is 1:1.2 to 1:

50.

9. The device according to claim 8, It is characterized in that The ratio of the separator diameter or equivalent diameter to the lower reaction chamber diameter or equivalent diameter is 1:2 to 1:

10.

10. The device according to claim 1, It is characterized in that The height of the mixing section is 20% to 35% of the total height of the separator, the height of the separation section is 55% to 70% of the total height of the separator; and the height of the stabilizing section is 5% to 10% of the total height of the separator.

11. A method for hydrotreating Fischer-Tropsch oil using the device according to any one of claims 1 to 10, comprising the following steps: (1) The Fischer-Tropsch synthetic oil enters the separation chamber in the hydroprocessing unit, where the lighter part is carried away by the hydrogen flowing upward from the bottom and enters the hydrotreating reaction zone upward; the heavier part enters the hydrocracking reaction zone downward and undergoes a cracking reaction with the countercurrent hydrogen. The light components produced by the reaction separate upward from the hydrocracking reaction zone and enter the hydrotreating reaction zone, while the heavy components flow downward for reflux or are partially discarded and discharged as tail oil; (2) The hydrofinished stream flows upward into the separator disposed at the top of the hydrofinishing reaction zone, where it is separated into naphtha and diesel fractions. The separated uncracked heavy components flow downward into the hydrocracking reaction zone again for hydrocracking reaction; In step (1), the ratio of the catalyst bed equivalent diameter of each of the hydrotreating reaction zone and the hydrocracking reaction zone to the total height of each of the catalyst beds is 2:1 to 10:

1.

12. The method according to claim 11, It is characterized in that In step (1), the ratio of the catalyst bed equivalent diameter of each of the hydrotreating reaction zone and the hydrocracking reaction zone to the total height of each of the catalyst beds is 3:1 to 6:

1.

13. The method according to claim 11, It is characterized in that The properties of the Fischer-Tropsch synthetic oil are as follows: density 0.6 g / cm 3 ~1.0g / cm 3 ; Final distillation point 650℃~750℃.

14. The method according to claim 13, It is characterized in that The properties of the Fischer-Tropsch synthetic oil are as follows: density 0.7 g / cm 3 ~0.95g / cm 3 ; Final distillation point 680℃~720℃.

15. The method according to claim 11, It is characterized in that The operating conditions of the Fischer-Tropsch synthetic oil hydrogenation reaction process in step (1) are as follows: reaction temperature of 380°C to 450°C, reaction pressure of 3MPa to 15MPa, top reflux ratio of 1.2 to 4.5, hydrogen-to-oil volume ratio of 300 to 2000, liquid hourly volume space velocity of 0.1h -1 ~10.0h -1 .

16. The method according to claim 15, It is characterized in that The operating conditions of the Fischer-Tropsch synthetic oil hydrogenation reaction process in step (1) are as follows: reaction temperature of 400°C to 450°C, reaction pressure of 4MPa to 12MPa, top reflux ratio of 1.5 to 3.0, hydrogen-to-oil volume ratio of 400 to 1500, liquid hourly volume space velocity of 0.5h -1 ~10.0h -1 .

17. The method according to claim 11, It is characterized in that Open 2 to 3 side lines in the separation section of the separator to extract the required products.

18. The method according to claim 11, It is characterized in that At the top of the separator, the extraction temperature is 60℃~80℃, and the extracted components are gas and light naphtha fraction; the extraction temperature of the first side line is 150℃~190℃, and the component extracted from this side line is heavy naphtha fraction; the extraction temperature of the second side line is 300℃~380℃, and the component extracted from this side line is diesel fraction.

19. The method according to claim 18, It is characterized in that At the top of the separator, the first side line extraction temperature is 160°C to 180°C, and the component extracted from this side line is a heavy naphtha fraction; the second side line extraction temperature is 310°C to 370°C, and the component extracted from this side line is a diesel fraction.

Images