CN116064131A - Fischer-Tropsch synthesis oil hydrotreatment device and method - Google Patents
Fischer-Tropsch synthesis oil hydrotreatment device and method Download PDFInfo
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- C—CHEMISTRY; METALLURGY
- C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
- C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
- C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
- C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
- C10G65/12—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including cracking steps and other hydrotreatment steps
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P30/00—Technologies relating to oil refining and petrochemical industry
- Y02P30/20—Technologies relating to oil refining and petrochemical industry using bio-feedstock
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- Chemical & Material Sciences (AREA)
- Oil, Petroleum & Natural Gas (AREA)
- Engineering & Computer Science (AREA)
- Chemical Kinetics & Catalysis (AREA)
- General Chemical & Material Sciences (AREA)
- Organic Chemistry (AREA)
- Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
Abstract
The invention discloses a Fischer-Tropsch synthesis oil hydrotreatment device and a Fischer-Tropsch synthesis oil hydrotreatment method. The device comprises a separator, a reaction cavity and a heavy oil bin which are sequentially communicated from top to bottom; the reaction chamber is internally provided with: a hydrofining catalyst bed, a separation cavity, a hydrocracking catalyst bed and a hydrogen distribution cavity. According to the invention, the characteristics of large difference of hydrocarbon composition, olefin and oxygen content of each fraction of Fischer-Tropsch synthesis oil are fully considered, the Fischer-Tropsch synthesis oil with wide distillation range distribution is divided into a lighter part and a heavier part by arranging the separation cavity in the hydrotreatment device, the lighter part with high olefin and oxygen content is upwards fed into the hydrofining reaction zone to carry out olefin saturation, deoxidization and deacidification reactions, and the heavier part is directly downwards fed into the hydrocracking reaction zone to carry out hydrocracking reactions, so that the occurrence of excessive cracking of the lighter part is avoided, and the yield of high-quality fuel oil is ensured to the greatest extent.
Description
Technical Field
The invention belongs to the field of hydrogenation reaction, and particularly relates to a device and a method for hydrotreating Fischer-Tropsch synthetic oil.
Background
The Fischer-Tropsch synthesis reaction is a reaction for generating hydrocarbons by taking hydrogen and carbon monoxide as raw materials under the action of a synthesis catalyst. The Fischer-Tropsch synthesis raw material synthesis gas has wide sources and can be converted from coal, natural gas, coalbed methane, biomass and the like. The Fischer-Tropsch synthetic oil mainly comprises normal alkane, alkene and a certain amount of oxygen-containing compounds, has extremely low content of non-ideal components such as sulfur, nitrogen, aromatic hydrocarbon and the like, and has great difference from conventional mineral petroleum in terms of hydrocarbon composition and main properties. The Fischer-Tropsch synthesis oil heavy naphtha fraction is basically free of sulfur and nitrogen, has very low octane number (the most of olefin and alkane are straight chains), and has higher olefin content and oxygen content; the diesel fraction has extremely low sulfur, nitrogen and aromatic hydrocarbon content, high cetane number, intermediate olefin content and oxygen content, and low heavy oil olefin content and oxygen content. Therefore, fischer-Tropsch synthetic oil must be subjected to corresponding hydrogenation and upgrading to obtain petrochemical raw materials or transportation fuels meeting the use specifications.
The hydrogenation upgrading technology of Fischer-Tropsch synthetic oil reported in the industry is mainly based on the technology of obtaining middle distillate diesel oil and aviation kerosene. Such as CN104711019A, CN1854265A, CN1854266A, CN101177626A, US5689031, these processes generally include two parts, hydrofining and hydrocracking, and a fixed bed reactor is used, and the fischer-tropsch synthesis oil is generally obtained by hydrofining saturated olefins and removing oxygenates, then separating naphtha, diesel oil and heavy oil according to the distillation range, and then obtaining naphtha and high-quality diesel oil blending components through hydrocracking the heavy oil. The Fischer-Tropsch synthetic oil after hydrocracking has high n-alkane content in naphtha obtained by fractionation and high saturation, and is an ideal raw material for preparing olefin by steam cracking; the cetane number of the obtained diesel oil reaches 70-80, and the diesel oil does not contain sulfur, nitrogen and aromatic hydrocarbon, and is an ideal diesel oil component for vehicles. The process flows are complex, the hydrotreating and hydrocracking parts are completely separated, two fractionating systems are arranged in total, and the device investment and the energy consumption are high.
The Fischer-Tropsch synthesis oil is firstly divided into two or three different fractions by the patents of US5378348, US6309432, CN101230291A and the like, the heavy fraction enters a hydroisomerization cracking reactor for carrying out isomerism cracking reaction, then the heavy fraction and the light fraction are mixed and enter a hydrofining reactor for carrying out reaction, and the hydrogenation product enters a fractionating tower for fractionating. The process has the advantages that the load of the device is reduced, the process is simplified, but the invention adopts the traditional hydrogenation reaction fixed bed reaction zone, the height-diameter ratio (the ratio of the total height of the reactor bed layer to the diameter) is generally 2-10, the reaction temperature rise is higher, and particularly, the olefin content in Fischer-Tropsch synthetic oil is far higher than that of conventional mineral petroleum, so that the 'flying temperature' risk of the catalyst bed layer is increased. In order to reduce the temperature, a large amount of cold hydrogen is injected between each bed layer of the reactor to reduce the reaction temperature, the reactor has high energy consumption, and the cold hydrogen is difficult to bring excessive heat out of the reactor in time at the initial stage and during abnormal operation, so that the service life of the catalyst is reduced, and the catalyst can be disabled under serious conditions.
The height and diameter of the traditional hydrogenation reaction fixed bed are relatively high so as to ensure that gas-liquid materials are fully contacted with the solid catalyst, and the required reaction depth and efficiency are achieved. Dong Fangliang et al, in "one weight technology" 1998.1 (total 75), "determination of major structural parameters of hydrogenation reactor", mention that "the ratio of height to diameter of a conventional fixed bed reactor is 4 to 9% more" in order to avoid a small ratio of height to diameter "poor catalyst contact efficiency due to maldistribution of fluid". Patent CN109679689 a also mentions that the height-to-diameter ratio of the existing liquid phase hydrogenation reactor is generally 2.5-12. The design of the height to diameter ratio of the hydrogenation reactor bed becomes the solidification cognition of the person skilled in the art, a large number of industrial practice applications also prove that the design has rationality and more general adaptability, and the wide industrial success possibly leads the person to fail to research whether other better choices exist for different types of reactions more comprehensively and deeply, no related research report exists for a long time, or only research reports which prove that the existing height to diameter ratio is suitable for the design.
In addition, fischer-Tropsch synthetic oil mainly consists of straight-chain high-carbon alkane and alkene, especially the light fraction has higher heavy alkene content which can reach more than 50%, while the traditional petroleum-based catalytic cracking raw material mainly consists of polycyclic hydrocarbon containing side chains, and the molecular structures of the two are greatly different. Compared with polycyclic hydrocarbon molecules with side chains, the cracking activation energy of the long-chain hydrocarbon molecules is lower, and the cracking reaction is easier, so that cracked products can generate cracking reaction again in the cracking process, the reaction depth is uncontrollable, a large amount of gas and olefin are generated, and the coking and blocking phenomena of the reactor can be caused. However, the existing method can not eliminate the coking and blocking phenomenon of the traditional reactor. This is because the product cannot leave the system during the entire reactor bed flow of feedstock and hydrogen, and the material flows downward as it reacts during the bed advancement. During this period, both fresh material reacts under the action of the catalyst and cracked products undergo cracking again.
Disclosure of Invention
Aiming at the defects of the prior art, the invention aims to provide a Fischer-Tropsch synthesis oil hydrotreatment device and a Fischer-Tropsch synthesis oil hydrotreatment method. So as to effectively control the reaction depth, greatly improve the yield of heavy naphtha and diesel oil, eliminate the risk of bed layer 'flying temperature', and prolong the service life of the catalyst.
In a first aspect the present invention provides a Fischer-Tropsch synthesis oil hydrotreater apparatus, the apparatus comprising:
(1) The separator, the reaction cavity and the heavy oil bin are sequentially communicated from top to bottom;
(2) The reaction chamber is internally provided with: a hydrofining catalyst bed, a separation cavity, a hydrocracking catalyst bed and a hydrogen distribution cavity.
In the above technical scheme, the ratio of the equivalent diameter of the cross-sectional area of the reaction chamber (the equivalent diameter formula is de=4a/L, a is the cross-sectional area of the reaction chamber, and the perimeter of the L bed layer) to the height is 2:1-10:1, preferably 3:1-6:1.
In the technical scheme, the height of the reaction cavity is generally 100-5000 mm, preferably 200-1000 mm. At higher aspect ratios, the flux of material through the catalyst bed located in the reaction chamber can be greatly increased while reducing the residence time of material and heat within the catalyst bed.
In the above technical solution, the cavity of the reaction chamber may be in a horizontal tank or a cylinder, and preferably in a horizontal tank. The two sides of the cavity are provided with sealing heads, so that the catalyst can be conveniently loaded and unloaded. The reaction cavity is filled with catalyst to form a reaction zone, and the reaction zone is separated into a plurality of reaction zone units by a net partition.
In the technical scheme, the height of the separation cavity arranged in the reaction cavity is generally 10-30% of the height of the reaction cavity. The ratio of the height of the hydrofining catalyst bed to the height of the hydrocracking catalyst bed is 5:1-1:8, preferably 2:1-1:6.
In the above technical scheme, the separation chamber is internally provided with a liquid distribution assembly, and the liquid distribution assembly comprises a liquid distributor, and a liquid distribution disc and a distribution cone which are arranged above the liquid distributor. For dispersing the liquid produced by the hydrogenation reactor into small droplets, the lighter fraction being carried upward under the hydrogen stripping action, the heavier fraction entering the hydrocracking reaction zone. The lighter fractions are typically naphtha and diesel fractions, and the heavier fractions are typically heavy diesel and wax oil fractions.
In the above technical solution, the liquid distributor is a conventional distributor in the field, such as a shower nozzle distributor, a coil pipe distributor, a porous straight pipe distributor, a straight pipe baffle type distributor, a baffle plate type distributor, a tangential circulation type distributor, a rotating vane distributor, a double vane type distributor, and the like. In the invention, the liquid distributor is preferably a porous tubular distributor and a straight pipe baffle type distributor, and the diameter of the pore canal of the tubular distributor is 0.5-20 mm, preferably 2-10 mm. The farther from the feed oil inlet end, the larger the pore size. The height of the liquid distributor from the top of the reactor bed is 1-1000 mm, preferably 50-500 mm. The height is related to the nature of the raw oil, the temperature and the pressure. Generally, the higher the temperature, the higher the height of the liquid distributor from the bed, so that the distributor can more uniformly drop the feedstock to the surface of the bed in a higher space. Also, the higher the pressure, the greater the spray angle of the liquid distributor, and the lower the height from the top of the reactor bed, the more space-saving.
In the technical scheme, the shape of the liquid distribution disc is the same as the cross section of the catalyst bed of the hydrogenation main reactor, and the area of the liquid distribution disc is 10-100%, preferably 60-100%, of the cross section area of the catalyst bed.
In the technical scheme, a plurality of first through holes are uniformly formed in the liquid distribution plate, first overflow rings are arranged around the first through holes, and overflow parts are arranged on the outer edge of the distribution plate; the aperture ratio of the distribution plate is 5-90%, the diameter of the first through hole is 5-100 mm, and the height of the first overflow ring is 1-30 mm.
In the technical scheme, the inner side of the first overflow ring is provided with the sawtooth part, the sawtooth part is bent downwards, and the sawtooth part is provided with the diversion trench.
In the above technical scheme, the distribution cone is arranged at the center of the upper part of the liquid distribution disc, the distribution cone is provided with a plurality of second through holes, and second overflow rings are arranged around the second through holes. The vertex angle of the distribution cone is larger than 90 degrees, the aperture ratio of the distribution cone is 5-80%, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2% -15% of the area of the liquid distribution disc.
In the above technical scheme, the ratio of the diameter or equivalent diameter of the separator to the diameter or equivalent diameter of the lower reaction chamber is 1:1.2-1:50, preferably 1:2-1:10. The diameter of the upper separator 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 the complete replacement of the fractionating tower is realized.
In the technical scheme, the separator sequentially comprises a mixing section, a separating section and a stabilizing section from bottom to top, wherein the height of the mixing section is 20% -35% of the total height of the separator, and the height of the separating section is 55% -70% of the total height of the separator; the height of the stabilizing section is 5-10% of the total height of the separator.
In the above technical scheme, the separation section is provided with a packing or a column plate. The packing or the tower plate is in a conventional form in the field, for example, the packing can be one or more random packing materials such as pall rings, raschig rings, saddle-shaped, open pore ring types, semi-rings, ladder rings, double arcs, halfpace rings, conjugate rings, flat rings, flower rings and the like, and the packing can also be metal or ceramic corrugated packing materials. The tray is one or more of bubble plate, sieve plate, floating valve plate, mesh plate, tongue plate, guide sieve plate, multi-downcomer tray, etc., or is tray without downcomer, such as flow-through sieve plate, flow-through corrugated plate, etc. Preferably to high efficiency trays such as float valves, sieve trays, and the like. The mixing section and the stabilizing section are not limited to whether or not to place the filler, and the reaction zone may be increased according to the process requirements.
In the above technical scheme, a plurality of separators are arranged in parallel in the hydrogen distribution cavity and the catalyst bed layer in the reaction cavity along the vertical direction, the hydrogen distribution cavity is divided into a plurality of air inlet units by the plurality of separators, and at least one hydrogen inlet is arranged at the bottom of each air inlet unit. A plurality of holes are distributed on each baffle; the separator extends upwards to the catalyst layer, the aperture ratio of the separator below the catalyst bed layer is less than 70%, and the aperture ratio of the separator in the catalyst layer is more than 50%.
In the technical scheme, the hydrogen inlet is provided with the gas distributor. In the invention, the gas distributor is preferably a tangential circulation type distributor and a rotary blade distributor, and the gas distributor can ensure that the flow speed of the gas entering the interface of the whole catalyst bed is relatively uniform, thereby avoiding the conditions of bias flow, channeling and the like.
The hydrogen distribution cavity of the hydrotreater is provided with the hydrogen feeding pipes, and the hydrogen feeding pipes are provided with a plurality of inlets, each hydrogen feeding pipe corresponds to a catalyst bed area between two partition boards, so that hydrogen can upwards pass through a reaction area at the top after coming out of each distributor. At least one pore canal is arranged at the joint of the baffle plate with holes and the bottom of the main reactor.
In the above technical scheme, the bottom of the catalyst bed between every two adjacent separators corresponds to 1-3 gas distributors, and the distribution area of the hydrogen gas from all the gas distributors in the separator area when reaching the bottom of the corresponding area bed should cover the bottom of the whole area bed. Further, the partition plate is annular or in a circular shape.
In the technical scheme, the heavy oil bin is arranged at the center of the bottom of the reaction cavity and is communicated with the plurality of air inlet units.
In the above technical solution, the hydrotreating apparatus further includes: one end of the reboiler is connected with an outlet of the heavy oil bin, and the other end of the reboiler is connected with the hydrogen distribution cavity. The temperature of the heavy oil sump is maintained at the temperature required by the reaction bed by means of a reboiler.
In the above technical solution, the hydrotreating apparatus further includes: and the liquid raw material inlet of each stage of auxiliary reaction cavity is connected with the heavy oil bin of the previous stage, and the top of the multistage auxiliary reaction cavity is connected to the separator.
The second aspect of the invention provides a Fischer-Tropsch synthesis oil hydrotreating method, which comprises the following steps:
(1) The Fischer-Tropsch synthesis oil enters a separation cavity in the hydrotreater, wherein a lighter part is carried away upwards by hydrogen flowing upwards from the bottom and enters a hydrofining reaction zone; the heavier part enters a hydrocracking reaction zone downwards to carry out a cracking reaction with countercurrent upward hydrogen, and the light component generated by the reaction is separated upwards from the hydrocracking reaction zone and then enters a hydrofining reaction zone, and the heavy component flows downwards to be refluxed or partially thrown outwards to be discharged as tail oil;
(2) The material flow after hydrofining upwards enters a separator arranged at the upper part of the hydrofining reaction zone, naphtha and diesel oil fractions are obtained through separation, and the separated uncracked heavy components downwards enter the hydrocracking reaction zone again for hydrocracking reaction.
In the above technical solution, after the fischer-tropsch synthesis oil in step (1) enters the separation chamber, it is preferably dispersed into a lighter fraction and a heavier fraction by flash evaporation and liquid distribution components.
In the above technical scheme, the ratio of the equivalent diameter of the catalyst bed (where the equivalent diameter formula is de=4a/L, a is the bed cross-sectional area, and L bed perimeter) to the total height of the catalyst beds in the hydrofining reaction zone and the hydrocracking reaction zone in the step (1) is 2:1 to 10:1, preferably 3:1 to 6:1.
The equivalent diameter to total catalyst bed height ratio refers to the sum of the catalyst bed heights when there are multiple catalyst beds. The cross-sectional area of the bed layer refers to the cross-sectional area of the reactor bed layer. The reactor bed is preferably of equal diameter, i.e. the cross-sectional area is the same at different locations throughout the catalyst bed. The cross-sectional area of the catalyst bed is generally the same as the cross-sectional area of the reaction chamber in the reactor, and cross-section refers to a top-down cross-section, i.e., a cross-section perpendicular to a vertical line within the reaction chamber. If there is a difference in the cross-sectional area of the reactor over the height of the catalyst bed, the cross-sectional area here refers to the average of the cross-sectional areas of the catalyst bed or the cross-sectional area of the reaction chamber throughout the catalyst bed.
The Fischer-Tropsch synthesis oil comprises high-temperature Fischer-Tropsch synthesis full distillate oil and low-temperature Fischer-Tropsch synthesis full distillate oil, and preferably low-temperature Fischer-Tropsch synthesis full distillate oil. The Fischer-Tropsch synthetic oil has the following properties: density of 0.6g/cm 3 ～1.0g/cm 3 Preferably 0.7g/cm 3 ～0.95g/cm 3 The method comprises the steps of carrying out a first treatment on the surface of the The final distillation point is 650 ℃ to 750 ℃, preferably 680 ℃ to 720 ℃.
In the technical scheme, the hydrofining reaction zone in the step (1) is filled with hydrofining catalyst. The hydrofining catalyst is generally one, two or more selected from alumina or silicon-containing alumina as carrier and metals of group VIB and group VIII as active components, such as W, mo, co, ni.
In the technical scheme, the hydrocracking reaction zone in the step (1) is filled with a hydrocracking catalyst. The catalyst generally comprises an active component and a carrier, wherein the carrier component comprises one or more of alumina, silicon-containing alumina and molecular sieves, preferably molecular sieves, and the molecular sieves can be Y-type molecular sieves; the active component is one or more of VIB group and VIII group metals, wherein the VIB group metals are Mo and/or W, and the VIII group metals are Co and/or Ni.
The hydrofinishing catalyst and hydrocracking catalyst shapes may be any conventional existing hydrocracking catalyst shape, preferably porous, heterogeneous and/or honeycomb catalysts. The pore diameter of the porous catalyst is 1-50 mm, preferably 4-20 mm; the average particle diameter of the shaped catalyst is 2-50 mm, preferably 4-30 mm; the diameter or the side length of the honeycomb catalyst holes is 1-50 mm, preferably 3-15 mm; the void fraction of the catalyst bed is recommended to be 15-85%, preferably 20-75%.
In the technical scheme, the Fischer-Tropsch synthesis oil hydrogenation reaction process (comprising hydrofining and hydrocracking) in the step (1)) The operating conditions of (2) are as follows: the reaction temperature is 380-450 ℃, the reaction pressure is 3-15 MPa, the reflux ratio of the tower top is 1.2-4.5, the hydrogen oil volume ratio is 300-2000, and the liquid hourly space velocity is 0.1h -1 ～10.0h -1 . Preferred operating conditions are: the reaction temperature is 400-450 ℃, the reaction pressure is 4-12 MPa, the reflux ratio of the top of the tower is 1.5-3.0, the volume ratio of hydrogen to oil is 400-1500, and the liquid hourly space velocity is 0.5h -1 ～10.0h -1 。
In the above technical solution, the separator withdraws the product at the side line or at the top. The gas is drawn off from the top, and naphtha and diesel fractions are drawn off from the side stream.
In the above technical scheme, it is preferable to divide the separation section of the separator by 2-3 side lines for withdrawing the desired product. At the top of the separator, the temperature of extraction is 60-80 ℃, and the extracted components are gas and light naphtha fraction. The temperature of the 1 st side draw is 150 ℃ to 190 ℃, preferably 160 ℃ to 180 ℃, and the component extracted from the side draw is heavy naphtha fraction. The temperature of the 2 nd side draw is 300 ℃ to 380 ℃, preferably 310 ℃ to 370 ℃, and the component drawn from the side draw is diesel oil distillate.
Further, in the above technical solution, the side draw line may be provided with a reflux.
Through a great deal of research, for the gas-liquid-solid three-phase reaction process with the rapid decrease of the liquid phase quantity and the rapid increase of the gas phase quantity in the reaction, the gas phase quantity rapidly increases to occupy a great deal of bed gaps, so that the flow rate of the liquid phase is greatly increased. According to the conventional design, although the gas-liquid-solid three-phase contact is ensured to be sufficient, the effective reaction time of the liquid phase which needs to be further converted is reduced, the contact probability of the gas phase which does not need to be reacted again (such as the gas phase obtained by liquid phase conversion under the reaction condition) and the catalyst is increased, and for a system which needs more liquid phase conversion and gas phase control secondary reaction, the overall reaction effect is limited to a certain extent, and the reaction conversion rate, the selectivity and the like are generally difficult to further improve.
According to the research, when the overall airspeed is similar, aiming at the gas-liquid-solid three-phase hydrogenation reaction with the rapid decrease of the liquid phase and the rapid increase of the gas phase in the reaction process, when the countercurrent contact of hydrogen and raw oil gas is adopted, the diameter-to-height ratio of a catalyst bed layer in a reactor is obviously higher than that of the catalyst bed layer in the prior art, so that the generated gas phase rapidly leaves the catalyst bed layer, the adverse effect accumulation effect of the generated gas phase is small, the liquid phase can have more sufficient probability of reacting on the catalyst, the traditional recognition that the small height-to-diameter ratio can bring adverse effects such as bad contact effect is overcome, the effect of obviously improving the yield of a target product (the target product in the inferior diesel hydrogenation technology is heavy naphtha) is achieved, and the problems of easiness in flooding, limited hydrogen-oil ratio and the like of the countercurrent reactor are solved.
Compared with the prior art, the invention has the advantages that:
(1) The invention fully considers the characteristics of large difference of hydrocarbon composition, olefin and oxygen content of each fraction of Fischer-Tropsch synthetic oil, and separates Fischer-Tropsch synthetic oil with wide distillation range distribution into a lighter part and a heavier part by arranging a separation cavity in a hydrotreatment device, wherein the lighter part with high olefin and oxygen content is upwards fed into a hydrofining reaction zone to carry out olefin saturation and deoxidation deacidification reaction, and the heavier part is directly downwards fed into a hydrocracking reaction zone to carry out hydrocracking reaction, so that the dosage of hydrofining and hydrocracking catalysts is reduced, the occurrence of excessive cracking of the lighter part on the hydrocracking catalysts is avoided, and the yield of high-quality fuel oil is ensured to the greatest extent. Meanwhile, the light product generated after hydrocracking of the heavier part contains a small amount of olefin, and the olefin in the light product can be hydrogenated and saturated upwards through the hydrofining catalyst, so that the stability of the product quality is ensured.
(2) The design of the liquid distribution component can lead heavy component liquid flowing downwards in the separation cavity to be dispersed into proper small liquid drops, and the lighter part of the heavy component liquid can be directly carried out by means of the stripping action of hydrogen and does not enter the hydrocracking reaction zone to participate in the hydrocracking reaction. Meanwhile, the heavy components entering the hydrocracking reaction zone can be distributed more uniformly through the design of the liquid distribution component, and the problem of poor contact of reactants on the catalyst bed layer under the condition of large diameter-height ratio of the traditional reactor is solved. By adopting the reactor, the void ratio of the reactor can be smaller under the same technological conditions and product index requirements. The temperature rise control of the reaction bed is smaller and the allowable feed load is larger. The improvement of the pore structure can improve the flooding characteristic and ensure good mass transfer performance. Finally, the product properties have good controllability.
(3) By arranging the separator and performing flash evaporation and steam stripping, the invention can realize timely extraction of naphtha and diesel products, takes away a large amount of heat released by the reaction, controls the reaction degree, prevents excessive cracking and gasification of the naphtha and diesel products, and ensures the yield of the naphtha and diesel products to the maximum extent. Meanwhile, the partial pressure of the product is kept low all the time, so that the reaction speed is increased, the hidden danger of blocking the catalyst by byproducts is eliminated, the yield of the target product is improved, and the service life of the catalyst is prolonged.
Drawings
FIG. 1 is a schematic diagram of a Fischer-Tropsch synthesis oil hydrotreater according to the invention;
FIG. 2 is a schematic diagram of the hydrotreating process flow of Fischer-Tropsch synthetic oil and the schematic diagram of the hydrotreating apparatus in accordance with the present invention;
wherein: 1-Fischer-Tropsch synthetic oil; 2-hydrogen; 3-hot feed oil; 4-hot hydrogen; 5-hydrogenation reactor inlet line; 6-a reaction chamber of the hydrogenation reactor; 7, a heavy oil bin of the hydrogenation reactor; 8-a separation cavity of the hydrogenation reactor; 9-a mixed section isomerization pour point depressing catalyst bed; 10-a stabilizing section; 11-hydrofining catalyst bed; 12-a hydrocracking catalyst bed; 13-a separation section; 14-a liquid distribution assembly; 15-a mixing section; 16-separator; 17-separator overhead condenser; 18-a liquid separating tank; 19-hydrogen-rich and noncondensable gases; 20-heavy naphtha product; 21-diesel product; a 22-light naphtha fraction; 23-hydrogenating the heavy fraction; 24-refluxing the heavy fraction; 25-a circulating oil pump; 26-tail oil; 27-mesh baffles; 28-heating furnace.
Fig. 3 is a schematic side view of a separator plate in accordance with an embodiment of the present invention.
FIG. 4 is a schematic side view of a separator plate in the form of a circular segment, which is positioned only in the hydrogen distribution chamber, without the separator plate in the interior of the catalyst bed, according to another embodiment of the present invention.
Fig. 5 is a schematic top view of an annular separator according to an embodiment of the present invention.
FIG. 6 is a schematic side view of the annular partition according to two embodiments of the present invention, the partition of FIG. 6-2 being located only within the hydrogen distribution chamber, the catalyst bed being devoid of the annular partition.
Fig. 7 is a schematic top view of a liquid distribution assembly according to an embodiment of the present invention.
Fig. 8 is a schematic top and side view of a liquid distribution assembly according to another embodiment of the present invention.
Fig. 9 is a schematic top view of a first overflow ring according to an embodiment of the invention.
Fig. 10 is a schematic perspective view of the first overflow ring of fig. 8.
Detailed Description
The invention is further described below with reference to the accompanying drawings.
Fig. 1 and 2 show a fischer-tropsch synthesis oil hydrotreater and a process flow according to the invention. The Fischer-Tropsch synthesis oil 1 is sent into a heating furnace 28, heated to 380-450 ℃ to obtain hot raw oil 3, enters a hydrogenation reactor separation cavity 8 in a hydrogenation reactor reaction cavity 6 through a hydrogenation reactor inlet pipeline 5, is dispersed by a flash evaporation and liquid distribution assembly 14, and the lighter part is carried away by hydrogen flowing upwards from the bottom and enters a hydrofining catalyst bed 11 (comprising 11-1-11-6); the heavier fraction is sprayed down evenly over the hydrocracking catalyst bed 12 (including 12-1 to 12-6). Simultaneously, the hot hydrogen 4 heated by the heating furnace 28 of the hydrogen 2 is fed into the reaction chamber 6 of the hydrogenation reactor, uniformly moves upwards from the bottom of each reaction zone under the action of the reticular baffle 27 and the gas distributors 4-1 to 4-6, is in countercurrent contact with the heavier part sprayed from the liquid distribution assembly in the hydrocracking catalyst bed 12, and is in countercurrent contact with the lighter part flowing upwards from the separation chamber in the hydrofining catalyst bed 11. Under the pressure of 4-15 MPa, the heavy component of Fischer-Tropsch synthetic oil undergoes hydrocracking reaction on a hydrocracking catalyst bed, a part of long-chain molecules are broken into short-chain molecules, and polycyclic aromatic hydrocarbon is also broken partially; the lighter components of the Fischer-Tropsch synthetic oil, the light components of the cracking reaction and hydrogen are subjected to olefin hydrogenation saturation, hydrodeoxygenation, impurity removal and other reactions on the hydrofining catalyst bed layer 11.
The hydrocarbons or other gases of smaller molecules produced by the hydrogenation reaction are rapidly carried by the hydrogen gas to the separator 16, and a portion of the heavier fraction falls evenly to the surface of the hydrocracking catalyst bed 12 through the separation action of the mixing section 15 and the separation section 13, and downward through the hydrofining catalyst bed 11, the separation chamber 8 of the hydrogenation reactor and the liquid distribution assembly 14. The other part of the light components continue to move to the stabilizing section 10 at the top of the separator in the separator, flow out from the top of the separator 16, continue to be cooled by the condenser 17 at the top of the separator, and are discharged outside the hydrogen-rich and non-condensable gas 19 after being subjected to gas-liquid separation by the liquid separating tank 18, and the liquid hydrocarbon is refluxed or extracted to be used as a light naphtha fraction 22. Heavy naphtha product 20 is drawn out from 150-190 ℃ of the side line of the separator, and diesel product 21 is drawn out from 300-380 ℃ of the side line of the separator.
The heavy fraction 23 flows out from the heavy oil bin 7 of the hydrogenation reactor, is mixed with raw oil as reflux heavy fraction 24 through a circulating oil pump 25 and then is used as the feeding of the hydrogenation reactor, and can be partially used as tail oil to be discharged out of the device through the circulating oil pump 25.
Further, in one or more exemplary embodiments of the present invention, the reaction chamber 6 of the hydrogenation reactor may be a horizontal tank, as shown in fig. 1, which is disposed axially and laterally, and two ends of the horizontal tank are provided with sealing heads. Further, in one or more exemplary embodiments of the present invention, the reaction chamber 6 may also be an oblate cylinder can, which is disposed axially in the longitudinal direction.
Further, in one or more exemplary embodiments of the present invention, the shape of the partition matches the bottom of the reaction chamber 6, and when the reaction chamber 6 is a horizontal tank, the partition is a circular-cut partition, as shown in fig. 3 and 4; when the reaction chamber 6 is a flat cylindrical tank, the plurality of spacers are coaxial annular spacers, as shown in fig. 5 to 7. Further, in one or more exemplary embodiments of the present invention, a plurality of circular holes are distributed on each of the separators. Further, in one or more exemplary embodiments of the present invention, as shown in fig. 1, a plurality of mesh baffles 27 may extend up to the hydrofining catalyst bed 11, and the opening ratio of the separator at the bottom not in contact with the catalyst bed 11 is less than 70%, and the lower opening ratio is advantageous for increasing the resistance so that the hydrogen gas enters the hydrofining catalyst bed 11 as upward as possible, further functioning as a gas distributor. The aperture ratio of the partition plate in the hydrofining catalyst bed 11 is more than 50%, which is advantageous for more fully utilizing the catalyst.
Further, in one or more exemplary embodiments of the invention, the separator 16 includes, from bottom to top, the mixing section 15, the separation section 13, and the stabilization section 10.
Further, in one or more exemplary embodiments of the invention, the liquid distribution assembly 14 disposed at the bottom of the separation chamber 8 of the hydrogenation reactor comprises a liquid distribution tray and a distribution cone. The liquid distribution plate is arranged above the liquid distributor, the shape of the liquid distribution plate is the same as that of the top surface of the porous catalyst layer, a plurality of first through holes are uniformly formed in the liquid distribution plate, a first overflow ring is arranged around the first through holes, and overflow parts (not shown in the figure) are arranged on the outer edge of the distribution plate. The dispensing cone is arranged in the upper centre of the liquid dispensing disc, the dispensing cone being provided with a plurality of second through holes around which second overflow rings (not shown in the figures) are arranged.
Preferably, but not by way of limitation, in one or more exemplary embodiments of the present invention, as shown in connection with fig. 9 and 10, the inner side of the first overflow ring is provided with a serration 14-1, which is bent downward, and the serration is provided with a diversion trench 14-2. Illustratively, the channels open along the center of the serrations.
In one or more embodiments of the invention, the hydrogenation reactor further comprises an auxiliary reaction chamber. It should be appreciated that the auxiliary reaction chamber may be multi-stage. The auxiliary reaction chambers of each stage are independently supplied with hydrogen, the center of the bottom is independently provided with a heavy oil bin, the liquid raw material inlet of each auxiliary reaction chamber of each stage is connected with the heavy oil bin of the previous stage, and the tops of the auxiliary reaction chambers of the multiple stages are connected to the separator 16.
The invention will be further illustrated with reference to specific examples, but it should be understood that the scope of the invention is not limited by the specific embodiments.
Example 1
With the apparatus and process flow of the present invention shown in fig. 1 and 2, the raw oil is a fischer-tropsch synthetic oil, and the properties are shown in table 1. After raw oil and hydrogen are heated to 430 ℃ by a heating furnace, the raw oil and the hydrogen enter a hydrogenation reactor, a hydrofining catalyst FHUDS-5 bed layer and a hydrocracking catalyst FC-14 bed layer produced by China petrochemical and Dalian petrochemical institute are sequentially placed in the hydrogenation reactor from top to bottom, and the equivalent diameter ratio of each catalyst bed layer is 4:1, the height of the hydrofining catalyst bed and the height of the hydrocracking catalyst bed are respectively 800mm. The height of the separation cavity between the hydrofining catalyst bed and the hydrocracking catalyst bed is 15% of the height of the reaction cavity. Annular baffles are arranged in the catalyst bed layer, and the number of the baffles is 4. A plurality of holes are distributed on the partition board; the separator extends upward to the catalyst layer, the aperture ratio of the separator below the catalyst layer is 40%, and the aperture ratio of the separator within the catalyst layer is 70%.
The raw material liquid is firstly subjected to light and heavy component separation through a liquid distribution assembly, the light component is hydrofined upwards and then enters a separator, the heavy component enters 4 different sub-catalyst bed areas in a hydrocracking catalyst bed, and the heavy component and hydrogen entering from the bottom react with hydrocracking catalyst, olefin saturation and the like. The light fraction produced by the reaction is quickly separated from the reaction system upwards, enters a separator after hydrofining, and the separator sequentially comprises a mixing section, a separating section and a stabilizing section from bottom to top, wherein the height of the mixing section is 30% of the total height of the separator, and the height of the separating section is 65% of the total height of the separator; the height of the stabilizing section is 5% of the total height of the separator. The mixing section and the stabilizing section are provided with packing raschig rings. The stream after hydrofining is mixed with light components at the lower part of a mixing section of a separator, then continues to upwards, the heavy fraction downwards enters a hydrocracking catalyst bed again for cracking reaction under the action of separation trays of the mixing section and a separation section, the light fraction is extracted as heavy naphtha fraction and diesel fraction at a lateral line and is sent out as products, noncondensable gas continues to be cooled by a condenser, after gas-liquid separation of a liquid separating tank, liquid hydrocarbon is extracted as light naphtha fraction, and gas enters desulfurization and deamination equipment for purification and recycling. Heavy fraction which is not cracked enough flows out from the bottom of the heavy oil bin, passes through the circulating pump and then enters the inlet of the hydrotreater to be used as circulating oil, the specific operation process conditions are shown in table 2, and the product distribution and properties are shown in table 4.
The liquid distribution assembly comprises a liquid distributor, a liquid distribution disc and a distribution cone, wherein the liquid distribution disc and the distribution cone are 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 uniformly provided with a plurality of first through holes, first overflow rings are arranged around the first through holes, the outer edge of the distribution plate is provided with overflow parts, the opening rate of the liquid distribution plate is 50%, the diameter of each first through hole is 10mm, and the height of each first overflow ring is 10mm. The distribution cone is arranged at the center of the upper part of the liquid distribution disc, a plurality of second through holes are formed in the distribution cone, and second overflow rings are arranged around the second through holes; the vertex angle of the distribution cone is 120 degrees, the aperture ratio of the distribution cone is 50 percent, and the height of the second overflow ring is 10mm; the bottom area of the dispensing cone is 10% of the area of the liquid dispensing tray.
Example 2
This example differs from example 1 in that the equivalent diameter height ratio of the hydrofining and hydrocracking catalyst beds is 5:1, the number of the separators in the catalyst bed is 6. A plurality of holes are distributed on the partition board; the separator extends upward to the catalyst layer, the aperture ratio of the separator below the catalyst layer is 30%, and the aperture ratio of the separator within the catalyst layer is 80%. The area of the liquid distribution tray was 90% of the cross section of the catalyst bed. The aperture ratio 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 center of the upper part of the liquid distribution disc, a plurality of second through holes are formed in the distribution cone, and second overflow rings are arranged around the second through holes; the vertex angle of the distribution cone is 150 degrees, the aperture ratio of the distribution cone is 70 percent, and the height of the second overflow ring is 20mm; the bottom area of the dispensing cone is 15% of the area of the liquid dispensing tray. The other conditions were the same as in example 1.
Example 3
This example differs from example 1 in that the hydrofinishing and hydrocracking catalyst bed equivalent diameter to height ratio is 6:1, the number of the separators in the catalyst bed is 6. The area of the liquid distribution tray was 80% of the cross section of the catalyst bed. The aperture ratio of the liquid distribution plate is 70%, the diameter of the first through hole is 30mm, and the height of the first overflow ring is 20mm. The distribution cone is arranged at the center of the upper part of the liquid distribution disc, a plurality of second through holes are formed in the distribution cone, and second overflow rings are arranged around the second through holes; the vertex angle of the distribution cone is 140 degrees, the aperture ratio of the distribution cone is 60 percent, and the height of the second overflow ring is 20mm; the bottom area of the dispensing cone is 15% of the area of the liquid dispensing tray. The other conditions were the same as in example 1.
Example 4
This example differs from example 1 in that the height of the mixing section of the separator is 25% of the total height of the separator, and the height of the separation section is 65% of the total height of the separator; the height of the stabilizing section is 10% of the total height of the separator. The other conditions were the same as in example 1.
Comparative example 1
Adopts the conventional two-stage hydrogenation method, namely refining and cracking process. Both the refining reactor and the cracking reactor adopt a reaction process that raw materials and hydrogen flow in parallel flow from top to bottom. The Fischer-Tropsch synthesis oil is subjected to olefin saturation in a hydrofining reactor, after deoxidation and impurity removal reaction, hydrofining material flows enter a first separation and fractionation system to obtain refined high-pressure gas, naphtha, diesel oil and heavy oil, wherein the refined high-pressure gas flows back into the hydrofining reactor, the heavy oil enters a hydrocracking reactor to carry out hydrocracking reaction, after the reaction is finished, the heavy oil enters a second separation and fractionation system to obtain gas, naphtha and diesel oil fractions through separation, and all the separated uncracked heavy components flow back into the hydrocracking reactor to carry out hydrocracking reaction again. The catalyst used was the same as in example 1 and the process conditions are shown in Table 3.
TABLE 1 Main Properties of raw oil
| Project | Data |
| Density, g/cm 3 | 0.815 |
| Distillation range, DEG C | |
| Initial 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 |
Table 2 example process conditions
| Example 1 | |
| Reaction temperature, DEG C | 410 |
| Reaction pressure, MPa | 8.0 |
| Volumetric hydrogen to oil ratio | 1000：1 |
| Volume space velocity, h -1 | 1.5 |
Table 3 comparative example 1 process conditions
| Example 1 | |
| One-stage hydrofining reactor | |
| Reaction temperature, DEG C | 300 |
| Reaction pressure, MPa | 8.0 |
| Volumetric hydrogen to oil ratio | 300：1 |
| Volume space velocity, h -1 | 3.0 |
| Two stage hydrocracking | |
| Reaction temperature, DEG C | 410 |
| Reaction pressure, MPa | 8.0 |
| Volumetric hydrogen to oil ratio | 1000:1 |
| Volume space velocity, h -1 | 1.5 |
TABLE 4 product distribution and Properties
| Example 1 | Example 2 | Example 3 | Example 4 | Comparative example 1 | |
| Naphtha fraction | |||||
| Distillation range, DEG C | 65～160 | 65～160 | 65～160 | 65～160 | 65～160 |
| Yield, wt% | 32.2 | 31.3 | 33.8 | 32.3 | 18.4 |
| Density, g/cm 3 | 0.708 | 0.705 | 0.710 | 0.706 | 0.714 |
| Composition, wt% | |||||
| N-alkanes | 91.20 | 92.30 | 91.57 | 91.74 | 87.57 |
| Isoparaffin(s) | 6.54 | 6.52 | 6.81 | 7.02 | 10.24 |
| Diesel oil fraction | |||||
| Distillation rangeEnclosing, DEG C | 160～370 | 160～370 | 160～370 | 160～370 | 160～370 |
| Yield, wt% | 66.1 | 67.2 | 65.0 | 66.8 | 78.1 |
| Density, g/cm 3 | 0.781 | 0.773 | 0.765 | 0.780 | 0.793 |
| Condensation point, DEG C | -3 | -4 | -5 | -3 | -2 |
| Cetane number | 83 | 83 | 84 | 83 | 78 |
From the results in Table 4, the yield of diesel oil is significantly higher than that of comparative example 1, and the cetane number is more than 83, which is a good quality diesel oil blending component; the naphtha fraction mainly consists of alkane and is a raw material for preparing ethylene and propylene by high-quality steam cracking.
Example 5
The laboratory performed simulated calculations of the bed reaction temperature profiles of the examples and comparative examples using ansys version 19.0 software. The simulation conditions were as actual data inputs for examples and comparative examples. Simulation results show that the center temperature of the traditional fixed bed is highest, and the temperature change is normally distributed from the inlet end to the outlet end, so that the temperature of the reactor bed is relatively uniform. The simulated temperature rise change of the bed is shown in Table 5.
Table 5 bed simulated temperature rise variation
| Bed temperature point | Example 1 | Example 2 | Comparative example 1 |
| Maximum radial temperature difference, DEG C | 3.1 | 3.5 | 28.1 |
| Average temperature, DEG C | 411.2 | 411.5 | 431.6 |
As can be seen from the results of Table 5, the temperature difference of the catalyst beds in examples 1 to 2 according to the present invention was significantly lower than that in comparative example 1, the temperature difference was reduced from the conventional fixed bed 28.1℃to 3.5℃and the difference between the average temperature and the control temperature was small, indicating that the reactor according to the present invention has eliminated the overheating phenomenon of the hydrocracking reaction.
Claims (16)
1. A fischer-tropsch synthesis oil hydrotreater, the apparatus comprising:
(1) The separator, the reaction cavity and the heavy oil bin are sequentially communicated from top to bottom;
(2) The reaction chamber is internally provided with: a hydrofining catalyst bed, a separation cavity, a hydrocracking catalyst bed and a hydrogen distribution cavity.
2. The apparatus according to claim 1, wherein the ratio of reaction chamber cross-sectional area equivalent diameter to height is 2:1 to 10:1, preferably 3:1 to 6:1.
3. The device according to claim 1, characterized in that the reaction chamber has a height of 100mm to 5000mm, preferably 200mm to 1000mm.
4. The device according to claim 1, wherein the height of the separation chamber arranged in the reaction chamber is 10% -30% of the height of the reaction chamber; the ratio of the height of the hydrofining catalyst bed to the height of the hydrocracking catalyst bed is 5:1-1:8, preferably 2:1-1:6.
5. The apparatus of claim 1, wherein a liquid distribution assembly is disposed within the separation chamber, the liquid distribution assembly comprising a liquid distributor and a liquid distribution tray and cone disposed above the liquid distributor.
6. The apparatus of claim 5, wherein the liquid distribution tray has the same shape as the cross section of the catalyst bed of the hydrogenation main reactor, and the area of the liquid distribution tray is 10% to 100%, preferably 60% to 100%, of the cross section area of the catalyst bed.
7. The device according to claim 5, wherein the liquid distribution plate is uniformly provided with a plurality of first through holes, a first overflow ring is arranged around the first through holes, and an overflow part is arranged at the outer edge of the distribution plate; the aperture ratio of the distribution plate is 5-90%, the diameter of the first through hole is 5-100 mm, and the height of the first overflow ring is 1-30 mm.
8. The apparatus of claim 5, wherein the dispensing cone is disposed in the upper center of the liquid dispensing tray, the dispensing cone having a plurality of second through holes and a second overflow ring disposed around the second through holes. The vertex angle of the distribution cone is larger than 90 degrees, the aperture ratio of the distribution cone is 5-80%, and the height of the second overflow ring is 1-30 mm; the bottom area of the distribution cone is 2% -15% of the area of the liquid distribution disc.
9. The apparatus according to claim 1, wherein the ratio of 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.
10. The device according to claim 1, wherein the separator comprises a mixing section, a separating section and a stabilizing section from bottom to top, the height of the mixing section is 20% -35% of the total height of the separator, and the height of the separating section is 55% -70% of the total height of the separator; the height of the stabilizing section is 5-10% of the total height of the separator.
11. A process for the hydroprocessing of fischer-tropsch synthesis oil using the apparatus of any one of claims 1 to 10, comprising the steps of:
(1) The Fischer-Tropsch synthesis oil enters a separation cavity in the hydrotreater, wherein a lighter part is carried away upwards by hydrogen flowing upwards from the bottom and enters a hydrofining reaction zone; the heavier part enters a hydrocracking reaction zone downwards to carry out a cracking reaction with countercurrent upward hydrogen, and the light component generated by the reaction is separated upwards from the hydrocracking reaction zone and then enters a hydrofining reaction zone, and the heavy component flows downwards to be refluxed or partially thrown outwards to be discharged as tail oil;
(2) The material flow after hydrofining upwards enters a separator arranged at the upper part of the hydrofining reaction zone, naphtha and diesel oil fractions are obtained through separation, and the separated uncracked heavy components downwards enter the hydrocracking reaction zone again for hydrocracking reaction.
12. The process according to claim 11, wherein the ratio of the equivalent catalyst bed diameter of each of the hydrofinishing reaction zone and the hydrocracking reaction zone to the total height of the respective catalyst bed in step (1) is from 2:1 to 10:1, preferably from 3:1 to 6:1.
13. The method according to claim 11, characterized in that the fischer-tropsch synthesis oil properties are as follows: density of 0.6g/cm 3 ～1.0g/cm 3 Preferably 0.7g/cm 3 ～0.95g/cm 3 The method comprises the steps of carrying out a first treatment on the surface of the The final distillation point is 650 ℃ to 750 ℃, preferably 680 ℃ to 720 ℃.
14. The process of claim 11, wherein the operating conditions of the fischer-tropsch synthesis oil hydrogenation process of step (1) are as follows: the reaction temperature is 380-450 ℃, the reaction pressure is 3-15 MPa, the reflux ratio of the tower top is 1.2-4.5, the hydrogen oil volume ratio is 300-2000, and the liquid hourly space velocity is 0.1h -1 ～10.0h -1 The method comprises the steps of carrying out a first treatment on the surface of the Preferred operating conditions are: the reaction temperature is 400-450 ℃, the reaction pressure is 4-12 MPa, the reflux ratio of the top of the tower is 1.5-3.0, the volume ratio of hydrogen to oil is 400-1500, and the liquid hourly space velocity is 0.5h -1 ～10.0h -1 。
15. The process of claim 11, wherein the separator withdraws product at the side or top, gas at the top, naphtha and diesel fractions at the side; preferably 2 to 3 side lines are cut off in the separation section of the separator for withdrawing the desired product.
16. The process according to claim 15, wherein at the top of the separator, the temperature of extraction is 60 ℃ to 80 ℃, and the components extracted are gas and light naphtha fraction; the temperature of the 1 st side draw is 150 ℃ to 190 ℃, preferably 160 ℃ to 180 ℃, and the components drawn from the side draw are heavy naphtha fraction; the temperature of the 2 nd side draw is 300 ℃ to 380 ℃, preferably 310 ℃ to 370 ℃, and the component drawn from the side draw is diesel oil distillate.
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| CN102876371A (en) * | 2011-07-11 | 2013-01-16 | 中国石油化工股份有限公司 | Inferior raw material hydrocracking method |
| CN105778995A (en) * | 2016-04-18 | 2016-07-20 | 武汉凯迪工程技术研究总院有限公司 | Method and device for producing good-quality diesel oil through combined hydrogenation of low-temperature Fischer-Tropsch synthesis oil and inferior crude oil |
| CN112705120A (en) * | 2019-10-25 | 2021-04-27 | 中国石油化工股份有限公司 | Heavy oil processing device and processing method |
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| CN102876371A (en) * | 2011-07-11 | 2013-01-16 | 中国石油化工股份有限公司 | Inferior raw material hydrocracking method |
| CN105778995A (en) * | 2016-04-18 | 2016-07-20 | 武汉凯迪工程技术研究总院有限公司 | Method and device for producing good-quality diesel oil through combined hydrogenation of low-temperature Fischer-Tropsch synthesis oil and inferior crude oil |
| CN112705120A (en) * | 2019-10-25 | 2021-04-27 | 中国石油化工股份有限公司 | Heavy oil processing device and processing method |
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