CN116064148B - A two-stage hydrocracking method - Google Patents
A two-stage hydrocracking method Download PDFInfo
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- CN116064148B CN116064148B CN202111272113.3A CN202111272113A CN116064148B CN 116064148 B CN116064148 B CN 116064148B CN 202111272113 A CN202111272113 A CN 202111272113A CN 116064148 B CN116064148 B CN 116064148B
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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
- C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
- C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
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- 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
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/10—Process efficiency
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Abstract
本发明公开了一种两段加氢裂化方法，包括：原料油与氢气混合后进入第一段的加氢裂化预处理反应器发生预反应；然后气液分离的液相产物进入第二段的加氢裂化反应器进行第一加氢裂化反应；第一加氢裂化反应流出物进入第二段的催化反应蒸馏塔进行第二加氢裂化反应和精馏，得到相应产品。所述方法可有效解决反应产物及时从反应区分离的问题，提高了目标产物的收率，同时提高加氢裂化反应效率、降低装置能耗。
The invention discloses a two-stage hydrocracking method, comprising: the raw oil is mixed with hydrogen and then enters the first stage of the hydrocracking pretreatment reactor for pre-reaction; then the liquid phase product of gas-liquid separation enters the second stage of the hydrocracking reactor for the first hydrocracking reaction; the effluent of the first hydrocracking reaction enters the second stage of the catalytic reaction distillation tower for the second hydrocracking reaction and rectification to obtain the corresponding product. The method can effectively solve the problem of timely separation of the reaction product from the reaction zone, improve the yield of the target product, and at the same time improve the efficiency of the hydrocracking reaction and reduce the energy consumption of the device.
Description
Technical Field
The invention belongs to the field of hydrogenation reaction, and particularly relates to a two-stage hydrocracking method.
Background
The hydrocracking technology has the characteristics of strong raw material adaptability, high production operation and product scheme flexibility, good product quality and the like, can directly convert various heavy and inferior feeds into high-quality jet fuel, diesel oil and lubricating oil base materials which are urgently needed in the market, and the ethylene raw material prepared by steam cracking chemical naphtha and tail oil, and has become one of the most important heavy oil deep processing technologies in the modern oil refining and petrochemical industry, and is increasingly widely applied at home and abroad. Along with the increasing maturity of the hydrocracking technology in China, the energy consumption of the hydrocracking device is reduced, and the improvement of the product distribution gradually becomes a main target of the development of the hydrocracking technology.
The hydrocracking process can be divided into a first-stage serial hydrocracking process flow, a single-stage hydrocracking process flow and a two-stage hydrocracking process flow according to the processing flow.
CN100569923C discloses a two-stage hydrocracking process. VGO/or other heavy distillate oil is used as raw oil, and various high-quality clean products are obtained under the hydrocracking operation condition by adopting a two-stage process flow. The method can effectively reduce the operation severity of a pretreatment reaction zone, improve the treatment capacity of the device, prolong the service life of the catalyst and improve the product quality.
The two-stage hydrocracking technology disclosed in US10144884B2 is a two-stage full conversion process, and is mainly characterized by that it is an equipment called hydrocracking reactor container, and said equipment can be wholly divided into upper and lower portions, and its upper portion adopts a special component to implement gas-liquid separation, and its lower portion is a general trickle bed hydrogenation reactor. The device comprises an upper feed and a lower feed, wherein a material flow coming out of a front hydrocracking reactor enters an upper part for gas-liquid separation, a gas phase is separated from the top, a liquid phase falls into a hydrocracking reactor part, the lower feed is a mixture material flow of unconverted oil and circulating hydrogen, the material flow and the liquid phase falling from the upper separation part are mixed above a catalyst bed layer of the hydrocracking reactor, and the material flow falls through the catalyst bed layer for hydrocracking reaction. The effluent from the top and bottom of the reactor is separated by hot high fraction and cold high fraction, the hydrogen is recycled, the liquid phase separated by the separator enters a fractionating tower, the products are discharged from the top and side of the fractionating tower, and the tail oil at the bottom of the fractionating tower is recycled.
From the development history of hydrocracking, technological progress is mainly focused on the aspects of raw material change, process flow change, catalyst performance improvement and the like. However, the hydrocracking reactor is still a conventional and classical trickle bed reactor.
Disclosure of Invention
Aiming at the defects in the prior art, the invention aims to provide a two-stage hydrocracking method. The method can effectively solve the problem that the reaction products are timely separated from the reaction zone, improves the yield of target products, improves the hydrocracking reaction efficiency and reduces the energy consumption of the device.
The invention provides a two-stage hydrocracking method which comprises the following steps of adopting a two-stage process flow, mixing raw oil with hydrogen, then entering a first-stage hydrocracking pretreatment reactor for pre-reaction, then entering a second-stage hydrocracking reactor for first hydrocracking reaction through a liquid-phase product subjected to gas-liquid separation, and entering a second-stage catalytic reaction distillation tower for second hydrocracking reaction and rectification through a first hydrocracking reaction effluent to obtain a corresponding product.
In the technical scheme, the gas-liquid separation device is a high-pressure separator. The operation condition of the high-pressure separator is that the pressure is 2.0-20.0 MPa and the temperature is 40-260 ℃. The gas phase product obtained by gas-liquid separation is recycled, and is preferably recycled to the hydrocracking pretreatment reactor of the first stage. The circulating device is a circulating compressor.
In the technical scheme, the liquid phase product subjected to gas-liquid separation is subjected to heat exchange before entering the hydrocracking reactor of the second section, the heat exchange device is a heating furnace, the material flow after heat exchange is the feeding oil of the hydrocracking reactor, and the temperature after heat exchange is 300-400 ℃.
Further, in the above technical solution, the raw oil includes at least one of a diesel oil fraction, a vacuum distillate (VGO), and a mixed oil thereof (including a corresponding fraction of a secondary processing product). The distillation range of the raw oil is generally 140-600 ℃. The sulfur content in the raw oil is more than 0.05wt%, preferably 0.10wt% to 4.0wt%, and/or the nitrogen content is more than 50 mug/g, preferably 500 to 3000 mug/g, and/or the aromatic hydrocarbon content is more than 30wt%, preferably 60wt% to 90wt%.
In the technical scheme, the raw oil is mixed with hydrogen after heat exchange, and the temperature of the raw oil after heat exchange is 260-350 ℃.
In the technical scheme, at least one catalyst bed layer is arranged in the hydrocracking pretreatment reactor, preferably, the number of the catalyst bed layers is 1-4, and more preferably, a quenching hydrogen inlet is arranged between the catalyst bed layers.
Further, in the above technical scheme, the pre-reaction includes at least one of hydrodesulfurization, hydrodenitrogenation and hydro-saturation reactions.
In the technical scheme, the hydrogenation condition of the pre-reaction is that the total pressure is 2.0-20.0 MPa, the average reaction temperature is 330-400 ℃, the liquid hourly space velocity is 0.5-3.0 h -1, and the hydrogen-oil volume ratio is 200:1-2000:1.
Further, in the above technical scheme, the hydrocracking pretreatment reactor is filled with a catalyst. The catalyst is a hydrocracking pretreatment catalyst, which may be selected from commercial hydrocracking pretreatment catalysts, or may be prepared according to methods well known in the art. Commercial hydrocracking pretreatment catalysts preferably smooth selected FF series hydrocracking pretreatment catalysts developed by the petrochemical institute, such as FF-36.
Further, in the above technical scheme, the nitrogen content in the liquid phase of the pre-reaction effluent is 20-200. Mu.g/g, preferably 50-100. Mu.g/g.
In the above technical scheme, at least one catalyst bed is disposed in the hydrocracking reactor in the second stage, and preferably, the number of catalyst beds is 1-2.
In the technical scheme, the first hydrocracking reaction is a shallow hydrocracking reaction, and the cracking conversion rate of the raw oil is within 30wt%, preferably 10-30 wt%. The first hydrocracking reaction mainly cracks materials which are easy to react, is convenient for controlling the temperature of the lower catalytic distillation reactor, and simultaneously improves the feeding temperature of the catalytic distillation reactor.
In the technical scheme, the hydrogenation condition of the first hydrocracking reaction is that the total pressure is 2.0-20.0 MPa, the average reaction temperature is 280-400 ℃, the liquid hourly space velocity is 3.0-30.0 h -1, and the hydrogen-oil volume ratio is 200:1-2000:1.
Further, in the above technical scheme, the hydrocracking reactor of the second stage is filled with a catalyst. The catalyst is a hydrocracking catalyst. The catalyst may be selected from commercial hydrocracking catalysts, and may also be prepared according to methods well known in the art. Commercial hydrocracking catalysts are preferably FC-series hydrocracking catalysts developed by the petrochemical institute, such as FC-24, and the like.
Furthermore, in the technical scheme, the catalytic reaction distillation tower comprises a catalytic reaction unit and a rectification unit, wherein the rectification unit is arranged at the upper part of the catalytic reaction unit.
In the technical scheme, the liquid phase in the first hydrocracking reaction effluent enters a catalytic reaction unit of the catalytic reaction distillation tower to carry out a second hydrocracking reaction, and the gas phase in the first hydrocracking reaction effluent enters a rectification unit of the catalytic reaction distillation tower to carry out rectification.
In the above technical scheme, the gas phase in the stream after the second hydrocracking reaction enters the rectification unit, and the liquid phase in the stream flows out from the bottom of the catalytic reaction distillation column to obtain the corresponding product.
Further, in the above technical scheme, the rectification unit rectifies the gas phase generated by the catalytic reaction unit and the gas phase in the first hydrocracking reaction effluent.
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 liquid is adopted, the generated gas phase rapidly leaves a catalyst bed, the accumulation effect of adverse effects of the generated gas phase is small, the liquid phase can have more sufficient probability of reacting on the catalyst, the traditional recognition of adverse effects of poor contact effect and the like of the traditional reactor under a large height-to-diameter ratio is overcome, the effect of obviously improving the yield of a target product is obtained, and the problems of easiness in flooding, limited hydrogen-oil ratio and the like of the countercurrent reactor are solved.
In the technical scheme, the catalytic reaction unit further comprises a plurality of catalyst beds, preferably large-diameter-ratio catalyst beds, wherein the diameter-ratio is 2-800:1, preferably 5-500:1. The aspect ratio is the ratio of the cross-sectional equivalent diameter of the catalyst bed to the height of the catalyst bed. The catalyst beds are used for filling hydrocracking catalysts, the upper part of each catalyst bed is provided with an inclined surface, the liquid-phase feeding subunit is arranged above the topmost first catalyst bed, the liquid-phase feeding subunit is downwards guided to the first catalyst bed, the gas-phase feeding subunit is arranged between the catalyst bed of the upper layer and the inclined surface of the next layer, the gas-phase feeding of each layer upwards enters the catalyst bed, the gas-phase channel is in a relative isolation state with the gas-phase feeding subunit, and gas-phase products generated after the gas-phase feeding and the liquid-phase feeding react in the catalyst beds directly enter the gas-phase channel.
Further, in the above technical solution, the gas phase channel may be located at the outer side of the catalytic reaction unit and is annular and penetrates through all the catalyst beds from bottom to top. In addition, the gas phase passage may be provided in the middle of the catalytic reaction unit.
Further, in the above technical solution, the whole inclined surface may be designed as an inverted umbrella-shaped partition structure. The inverted umbrella baffle ends may be provided with an inner downcomer having a bottom spaced from the bottom of the catalyst bed by a distance such that the liquid phase feed enters the catalyst bed in a radial direction.
Further, in the technical scheme, the catalyst bed layer can be provided with an overflow weir which is arranged at one side close to the gas phase channel, and a liquid sealing baffle which is arranged at the upper part of the overflow weir and is used for isolating gas phase feeding materials from gas phase products.
Further, in the technical scheme, the liquid sealing baffle plate can comprise a horizontal part which is in an annular flat plate shape and is positioned above the overflow weir, and a vertical part which is in a cylindrical shape, wherein the vertical part and the horizontal part are integrally formed, and the lower end of the vertical part is separated from the bottom of the catalyst bed by a certain distance.
In the technical scheme, the liquid-phase feeding subunit further comprises a liquid-phase feeding pipe which extends along the radial direction of the catalytic reaction unit, and a liquid-phase distributing pipe which is annular and is orthogonal or tangential to the liquid-phase feeding pipe, wherein the pipe wall of the liquid-phase distributing pipe is provided with a plurality of pore canals for uniformly distributing liquid-phase feeding to all positions above the first catalyst bed layer.
In the technical scheme, the gas-phase feeding subunit can further comprise a gas-phase feeding pipe which extends along the radial direction of the catalytic reaction unit, and a gas-phase distributing pipe which is annular or multi-layer concentric annular, wherein the gas-phase distributing pipe is orthogonal or tangentially intersected with the gas-phase feeding pipe, and a plurality of pore channels are arranged on the wall surface of the gas-phase distributing pipe and are used for uniformly distributing gas-phase feeding to all directions at the bottom of the catalyst bed.
In the technical scheme, the gas-phase feeding subunit can further comprise a gas-phase distributing disc which is positioned at the bottom of the catalyst bed layer and is in a disc shape as a whole, and a plurality of holes are uniformly distributed on the gas-phase distributing disc.
Further, in the above technical solution, the gas-phase distribution pipe may be disposed below the catalyst bed or within the catalyst bed.
Further, in the above technical scheme, the rectification unit may be provided with an inert packing layer, and the ascending gas phase product from the gas phase channel is rectified in the inert packing layer.
In the technical scheme, the inert filler layer is filled with inert filler, and the inert filler can be selected conventionally according to the requirement, and is preferably at least one of Raschig rings and porcelain balls.
In the technical scheme, a flow guide subunit is arranged between the inert filler layer and the first catalyst bed layer, the flow guide subunit is of an inverted umbrella-shaped baffle plate structure as a whole, a hollow structure is arranged at a position, corresponding to the gas phase channel, of the baffle plate, which is close to one side of the distillation tower wall, and one side, close to the middle part of the distillation tower, of the baffle plate is provided with an opening for allowing liquid-phase heavy components falling after rectification to enter the catalytic reaction unit.
Further, in the above technical scheme, the catalytic reaction distillation column may be a multi-layer plate column structure. The number of catalyst beds in the catalytic reaction distillation tower is more than 2, preferably 2-20.
Further, in the above technical scheme, the catalytic reaction distillation tower is suitable for a reaction system in which at least one liquid-phase feed and at least one gas-phase feed are subjected to chemical reaction on a hydrocracking catalyst, and at least one gas-phase product in reaction products is formed.
In the technical scheme, the second hydrocracking reaction condition of the catalytic reaction unit is that the average reaction temperature is 300-450 ℃, the liquid hourly space velocity is 0.5-3.0 h -1, and the hydrogen oil volume ratio is 500:1-10000:1.
In the above technical scheme, the catalyst of the hydrocracking reaction used by the catalytic reaction unit is a hydrocracking catalyst. The catalyst may be selected from commercial hydrocracking catalysts, and may also be prepared according to methods well known in the art. Commercial hydrocracking catalysts are preferably FC-series hydrocracking catalysts developed by the petrochemical institute, such as FC-24, and the like.
In the technical scheme, the rectification reaction condition of the rectification unit is that the pressure at the top of the tower is 2.0-20.0 MPa and the temperature at the top of the tower is 200-400 ℃.
In the technical scheme, the reaction effluent of the rectifying unit is subjected to gas-liquid separation after heat exchange and cooling, one part of the obtained liquid phase material flow flows back to the top of the rectifying unit of the distillation tower, and the other part of the obtained liquid phase material flow is the obtained light product, such as naphtha. The weight ratio of the portion of the reflux to the distillation column to the product portion of the effluent line, referred to as the overhead reflux ratio, may be determined based on product quality requirements. The reflux ratio of the tower top is 1:1-3:1.
Further, in the above-described technical scheme, heavy products such as diesel oil fraction, hydrocracking tail oil and the like flow out from the bottom of the catalytic reaction distillation column.
The hydrocracking process of the present invention is a two-stage hydrocracking process. The hydrocracking raw material is mixed with hydrogen through heat exchange and then enters a hydrocracking pretreatment reactor to carry out hydrodesulfurization, hydrodenitrogenation, hydrogenation saturation and other reactions, after gas-liquid separation, the obtained liquid phase material flow enters the hydrocracking reactor to carry out first hydrocracking reaction, namely shallow hydrocracking reaction, and the effluent of the first hydrocracking reaction directly enters a catalytic reaction distillation tower. The catalytic reaction distillation tower is divided into an upper section and a lower section, wherein the upper section is a rectifying section, and the lower section is a reaction section. The gas phase in the effluent of the first hydrocracking reaction enters the rectifying section upwards, the liquid phase enters the reaction section downwards, and the second hydrocracking reaction, namely the main hydrocracking reaction, is carried out through the catalyst bed layer. The reaction section is provided with a plurality of catalyst beds, the upper surface of each catalyst bed is connected with a gas phase channel, the gas phase enters the rectifying section through the gas phase channel, and the oil product which is still in liquid phase without reaction or reaction passes through the catalyst beds layer by layer, reaches the bottom of the catalytic reaction distillation tower and is still in liquid phase flows out of the distillation tower.
Compared with the prior art, the invention has the following beneficial effects:
(1) In the invention, the two-stage hydrocracking method adopts two-stage technological processes, comprising the steps of pre-reaction, first hydrocracking reaction of liquid-phase products after separation, second hydrocracking reaction and rectification in a catalytic reaction distillation tower. The method has stronger adaptability to raw materials, and can process VGO with high sulfur and nitrogen content and higher dry point and inferior straight run or secondary processed oil products doped with DAO. In addition, the process of the invention can effectively reduce the temperature of the hydrocracking reactor and improve the hydrocracking airspeed. And the load of the rectifying section of the catalytic distillation tower is reduced to a certain extent, and the corrosion of the heat exchange condenser at the top of the tower is obviously reduced.
In the invention, the catalytic reaction distillation tower is used, and the gas-phase product generated after the hydrocracking reaction of the reactant in the catalyst bed can timely leave the reaction zone, and can not enter the upper catalyst bed again, so that the secondary reaction of the target gas-phase product is avoided, and the selectivity of the target product is improved. The gas phase product in the reaction zone is separated in time, so that the reaction driving force is increased, the equilibrium conversion rate is improved, the yield of the target product is high, the yield of the components below C4 is low, the hydrogen consumption is effectively reduced, and the aromatic hydrocarbon content of the product at the top of the catalytic reaction distillation tower for the hydrocracking reaction is obviously increased.
(2) The invention carries out intensive research on hydrogenation reaction with rapid decrease of liquid phase quantity and rapid increase of gas phase quantity under reaction conditions, breaks through the long-term design scheme in the field, provides a technical means with completely different design parameters from the prior art, provides a proper technical scheme such as a process flow, a catalyst bed structure and the like, achieves outstanding technical effects, exceeds the expectations of the technicians in the field, and breaks through the routine solidification cognition in the field.
The invention can lead lighter parts to be directly carried out by arranging a plurality of fixed bed layers (particularly a large diameter-height ratio bed layer) in the catalytic reaction unit of the catalytic reaction distillation tower, 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 phase distribution pipe, and the problem of poor contact of reactants on a catalyst bed layer under a large height-diameter 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.
In addition, the invention ensures that the light component of the distillate oil after hydrogenation in the second hydrocracking reaction can be quickly separated from the system to not excessively participate in the cracking reaction by reasonably setting the reaction flow and controlling the catalyst bed structure, and the product can be quickly separated from the reaction system, so that 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 while the positive reaction speed is increased.
The invention can realize the timely extraction of light intermediate products through flash evaporation and steam stripping by arranging the rectification unit of the catalytic reaction distillation tower so as to effectively control the reaction degree, furthest reserve aromatic components and become good chemical raw materials. Meanwhile, the partial pressure of the product is kept low all the time, so that the reaction speed is increased, the reaction efficiency is improved, and bad components such as hydrogen sulfide, ammonia and the like which are easy to coke can be taken away.
Drawings
FIG. 1 is a schematic structural view of a catalytic reaction distillation column according to the present invention.
Fig. 2 is a top view of a liquid distribution pipe in the catalytic reaction unit of the present invention.
FIG. 3 is a top view of the gas phase feed pipe and gas phase distribution pipe in the catalytic reaction unit of the present invention (showing the case where the gas phase feed pipe is disposed orthogonally to the annular gas phase distribution pipe).
FIG. 4 is another top view of the gas phase feed tube and gas phase distribution tube in the catalytic reaction unit of the present invention (showing the tangential intersection of the gas phase feed tube and annular gas phase distribution tube).
FIG. 5 is a top view of a gas distribution tube of the present invention employing concentric dual ring distribution tubes.
FIG. 6 is a top view of a gas phase distributor plate in a catalytic reaction unit of the present invention.
FIG. 7 is a top view of a catalyst support plate in a catalytic reaction unit of the present invention.
FIG. 8 is a top view of an inert packing support tray in a rectification unit of the present invention.
Fig. 9 is a top view of a deflector unit in a rectification unit of the present invention.
FIG. 10 is a process flow diagram of the hydrocracking process of the present invention.
The main reference numerals illustrate:
The device comprises a 1-catalytic reaction distillation tower, a 102-hydrocracking pretreatment reactor, a 103-circulating compressor, a 104-high-pressure separator, a 105-hydrocracking reactor liquid phase feed heating furnace, a 106-hydrocracking reactor, a 107-heat exchange cooler, a 108-reflux tank and a 109-circulating hydrogen heating furnace;
1001-raw oil, 1002-raw oil and hydrogen gas mixture, 1003-pre-reaction effluent, 1004-hydrotreatment stage high-pressure separator gas phase effluent, 1005-hydrotreatment stage recycle gas, 1006-hydrotreatment stage make-up hydrogen, 1007-hydrotreatment stage high-pressure separator liquid phase effluent, 1008-hydrocracking reactor feed oil, 1009-hydrocracking reactor mixed feed, 1010-hydrocracking reactor effluent, 1011-catalytic distillation column overhead effluent, 1012-reflux tank feed, 1013-reflux tank overhead gas, 1014-hydrocracking reactor recycle gas, 1015-catalytic distillation column recycle gas, 1016-catalytic distillation column feed hydrogen, 1017-reflux tank bottom liquid phase stream, 1018-catalytic distillation column overhead reflux, 1019-catalytic distillation column overhead product (light product), 1020-catalytic distillation column bottom product (heavy product), 1021-make-up hydrogen gas.
10-Hydrocracking catalyst, 11-first umbrella-shaped partition plate, 12-inner downcomer, 13-gas phase channel, 14-overflow weir, 15-liquid falling folded plate, 16-liquid receiving disc, 17-liquid sealing baffle, 18-outer downcomer, 19-catalyst supporting disc and 191-grille;
21-a liquid phase feeding pipe, 22-a liquid phase distributing pipe, 220-a liquid phase distributing pipe body and 221-a liquid phase pore canal;
31-gas phase feeding pipe, 32-gas phase distributing pipe, 320-gas phase distributing pipe body, 321-gas phase pore canal, 33-gas phase distributing disk and 331-pore;
41-second umbrella-shaped partition plates, 411-hollowed parts, 42-liquid phase heavy component collecting ports, 43-filler supporting plates and 44-inert fillers.
Detailed Description
The present invention is further illustrated by the following examples, but the scope of the present invention is not limited by the examples.
Throughout the specification and claims, unless explicitly stated otherwise, the term "comprise" or variations thereof such as "comprises" or "comprising", etc. will be understood to include the stated element or component without excluding other elements or other components.
Spatially relative terms, such as "below," "beneath," "lower," "above," "upper," and the like, may be used herein for ease of description to describe one element's or feature's relationship to another element's or feature's in the figures. It will be understood that the spatially relative terms are intended to encompass different orientations of the article in use or operation in addition to the orientation depicted in the figures. For example, if the article in the figures is turned over, elements described as "below" or "beneath" other elements or features would then be oriented "above" the elements or features. Thus, the exemplary term "below" may encompass both a direction of below and a direction of above. The article may have other orientations (rotated 90 degrees or other orientations) and the spatially relative descriptors used herein interpreted accordingly.
The terms "first," "second," and the like herein are used for distinguishing between two different elements or regions and are not intended to limit a particular position or relative relationship. In other words, in some embodiments, the terms "first," "second," etc. may also be interchanged with one another.
As shown in fig. 1, a catalytic reaction distillation tower 1 of the present invention has an up-down structure, a catalytic reaction unit is disposed at the lower part of the tower, and a rectification unit is disposed at the upper part of the tower. The catalytic reaction unit is an internal component of the reactive distillation tower 1 and comprises a plurality of catalyst beds, a liquid-phase feeding subunit, a gas-phase feeding subunit and a gas-phase channel. Wherein, the catalyst beds are used for filling the hydrocracking catalyst 10, and from the second layer at the top, the upper part of the hydrocracking catalyst 10 filled by each catalyst bed is provided with an inclined surface, and the overall shape formed by the inclined surfaces can be an inverted umbrella shape, which plays a role of a partition plate, on one hand, can separate gas phase feeding and product gas between adjacent bed layers, and on the other hand, plays a role of guiding liquid phase and gas phase, and the inverted umbrella surface can be an arc shape or a folded umbrella shape, which is preferable but not limited. The liquid phase feed subunit is disposed above the top-most first catalyst bed, and the liquid phase feed is led downward to the first catalyst bed to contact with the hydrocracking catalyst 10, specifically, the end of the inverted first umbrella-shaped partition 11 is provided with an inner downcomer 12 (i.e. a space surrounded by the liquid-lowering flaps 15 in fig. 1), and the bottom of the inner downcomer 12 is spaced from the bottom of the catalyst bed by a distance such that the liquid phase feed enters the catalyst bed along the radial direction of the reactive distillation column 1. The gas phase feeding subunit is arranged on each catalyst bed layer, specifically between the catalyst bed layer on the upper layer and the inverted first umbrella-shaped baffle plate 11 on the lower layer, and the gas phase feeding of each layer upwards enters the catalyst bed layer. After the gas-liquid phase feed and the hydrocracking catalyst 10 are fully reacted in the catalyst bed, the gas phase product of each layer is guided to the gas phase passage 13 along the lower portion of the inverted first umbrella-shaped partition 11. The gas phase channel 13 is in a relatively isolated state from the gas phase feeding subunit, i.e. the gas phase product generated after the gas phase feeding and the liquid phase feeding react in the catalyst bed directly enters the gas phase channel 13. Preferably, but not by way of limitation, the gas phase channels are located outside the reactive distillation column 1, are annular and run through all the catalyst beds from bottom to top.
As further shown in fig. 1 and 2, the liquid phase feed subunit further comprises a liquid phase feed pipe 21 and a liquid phase distribution pipe 22. The liquid-phase feeding pipe 21 extends along the radial direction of the catalytic reaction unit, the liquid-phase distributing pipe 22 is annular, the liquid-phase feeding pipe 21 is orthogonal or tangentially intersected with the pipe body 220 of the liquid-phase distributing pipe 22, and the pipe wall of the liquid-phase distributing pipe 22 is provided with a plurality of pore canals 221 for uniformly distributing the liquid-phase feeding material to various positions above the first catalyst bed layer. The openings of the tunnel 221 may be in all directions above, below, and to the side faces of the tube. Liquid phase feed enters the reactive distillation column 1 through a liquid phase feed pipe 21, enters the column through an annular liquid phase distribution pipe 22 and is distributed to the first catalyst bed. The annular liquid distribution pipe 22 has an annular diameter larger than the outer diameter of the inner downcomer 12 and smaller than the inner diameter of the reactive distillation column 1. The height of the liquid-lowering folded plate 15 is generally smaller than the filling height of the catalyst layer, and the size of the enclosed space of the liquid-lowering folded plate 15 is determined according to the flow rate of the liquid-phase reactant of the layer.
As further shown in FIG. 1, the height of each catalyst bed in the reactive distillation column 1 can be the same or different, and the upper part of the catalyst bed is fixed by a screen according to different chemical reaction systems, so that the bed is relatively stable, and the height of the bed is set to be 10mm to 1000mm. The catalyst bed is provided with an overflow weir 14 and a liquid seal baffle 17, and the overflow weir 14 is arranged at one side close to the gas phase channel 13. A liquid seal 17 is provided above weir 14 to isolate the gas phase feed from the gas phase product. Further, the liquid seal baffle 17 comprises a horizontal part and a vertical part, the horizontal part is in a ring-shaped flat plate shape and is positioned above the overflow weir 14, the vertical part is in a cylinder shape, the vertical part and the horizontal part are integrally formed, other seamless connection modes can be adopted, and the lower end of the vertical part is separated from the bottom of the catalyst bed by a certain distance, so that the outflow of liquid phase products can be ensured. Unreacted liquid feed and reacted but liquid-phase-maintaining material within the catalyst bed passes over weir 14, through outer downcomer 18 (i.e., the annular space between weir 14 and the inner wall of gas phase channel 13), along inverted first umbrella baffle 11, through inner downcomer 12 of the next layer into the next catalyst bed. The height of weir 14 is above the upper level of the catalyst in the bed, preferably 10 to 100mm. The size of the space between the overflow weir 14 and the annular outer downcomer 18 formed by the inner wall of the gas phase channel 13 depends on the size of the liquid phase load, and the size of each bed downcomer can be the same or different.
As further shown in fig. 1,3 to 5, the gas phase feed subunit comprises a gas phase feed pipe 31 and a gas phase distribution pipe 32, the gas phase feed pipe 31 extending in the radial direction of the reactive distillation column 1. The gas distribution pipe 32 is annular (see fig. 3 and 4) or multi-layer concentric annular (see two-layer concentric rings of fig. 5), the gas inlet pipe 31 is orthogonal to the pipe body 320 of the gas distribution pipe 32 (see fig. 3) or tangentially intersects with the pipe body (see fig. 4), and a plurality of holes 321 are formed in the wall surface of the gas distribution pipe 32 and are used for uniformly distributing the gas phase feed to all directions of the bottom of the catalyst bed. Preferably, and not by way of limitation, the gas distribution tube 32 may be disposed below the catalyst bed or within the catalyst bed. As further shown in fig. 6, the gas phase feed subunit further includes a gas phase distribution plate 33, where the gas phase distribution plate 33 is located at the bottom of the catalyst bed and has a plate shape as a whole, and a plurality of holes 331 are uniformly distributed on the gas phase distribution plate. The gas phase feed enters the reactive distillation column 1 through a gas phase feed pipe 31 of each layer, is distributed into the reactive distillation column 1 through an annular gas phase distribution pipe 32, and enters the catalyst bed upward through a gas phase distribution plate 33 at the lower part of the catalyst support plate 19. The gas phase feeding pipe 31 enters the reactive distillation column 1 in a radial direction and is orthogonal or tangentially intersected with the annular gas phase distributing pipe 32, the annular gas phase distributing pipe 32 is positioned below the catalyst bed layer, the annular diameter of the annular gas phase distributing pipe 32 is smaller than the outer annular diameter of the catalyst bed layer, the inner diameter is larger than the inner annular diameter of the catalyst bed layer, and a plurality of pore canals 321 of the pipe wall of the annular gas phase distributing pipe 32 are convenient for gas to be uniformly distributed at all positions of the gas phase distributing disc 33. The function of the catalyst support plate 19 is mainly to support the catalyst bed, ensuring that the catalyst bed remains stable in the axial direction of the reactive distillation column. The purpose of the gas phase distributor plate 33 is to ensure uniform distribution of the gas phase feed while avoiding direct leakage of the liquid phase feed over the catalyst bed as much as possible (with the gas phase distributor plate 33 of the present invention, the liquid leakage is < 15%). On the same plane, when more than one concentric annular gas distribution tube 32 of different diameters is provided, the distribution of the gas phase feed may be made more uniform. The embodiment of fig. 1 provides for the annular gas distribution tube 32 to be positioned below the catalyst bed, and when the annular gas distribution tube 32 is installed within the catalyst bed, the catalyst support plate 19 may be modified from the grid 191 of fig. 7 to a support plate, while eliminating the gas distribution plate 33.
As further shown in fig. 1, after the liquid phase feed and the gas phase feed react in the catalyst beds of the respective layers, the produced gas phase product directly enters the annular gas phase channel 13 on the column wall side, the gas phase product passes through the gas phase channel 13 and rises up to the inert packing layer 44 of the rectifying unit, a grid-type packing support plate 43 (see fig. 8) is provided at the lower part of the inert packing layer 44, and the gas phase product from the catalytic reaction unit is rectified therein. The feed is converted into target products (namely gas-phase products) through the catalytic reaction unit, and is timely separated from a reaction system (the gas-phase feed is not mixed with the gas-phase products for secondary reaction), the gas-phase products obtained by the catalytic reaction unit rise to the rectification unit through the gas-phase channel 13, and one or two or more target products can be obtained through the rectification unit. In the rectification process of the inert filler layer 44, some liquid phase heavy components are generated, so that a flow guiding subunit is arranged between the inert filler layer 44 and the first catalyst bed layer, and is a connecting device between the rectification unit and the catalytic reaction unit, and the flow guiding subunit can play a role in uniformly distributing the gas phase product rising up from the catalytic reaction unit while collecting liquid. The flow guiding subunit is in an inverted umbrella-shaped baffle structure (namely, a second umbrella-shaped baffle 41 in fig. 1), the second umbrella-shaped baffle 41 is in a hollowed-out structure (see hollowed-out part 411 in fig. 9) at a position corresponding to the gas phase channel 13 on one side close to the distillation column wall, and the hollowed-out part 411 is arranged on one side close to the column wall so that ascending gas phase products from the gas phase channel 13 can smoothly enter the inert seasoning layer 44. The second umbrella-shaped partition 41 is provided with an opening (i.e., liquid-phase heavy component collecting opening 42 in fig. 1 and 9) for the liquid-phase heavy component falling down after rectification to enter the catalytic reaction unit on the side near the middle of the distillation column, and the liquid-phase heavy component falling down through the rectification unit can enter the catalytic reaction unit again.
In the catalytic reaction unit, liquid-phase feeding and gas-phase feeding are subjected to catalytic reaction in the catalyst bed, gas-phase products and unreacted gas-phase feeding are lifted to be separated from a reaction system through the gas-phase channel 13, and gas-phase products generated after chemical reaction of reactants in the catalyst bed can timely leave the reaction zone and cannot enter the upper catalyst bed again (isolated by the inverted umbrella-shaped partition plate), so that secondary reaction of target products is avoided, and the selectivity of the reaction is improved. Meanwhile, due to the fact that products in the reaction zone leave, the reaction driving force is increased, and the equilibrium conversion rate is improved.
The reactive distillation column of the present invention employs the above catalytic reaction unit, and the reactive distillation column 1 may be a multi-layer plate column structure. The number of the catalyst beds in the reactive distillation tower is two or more. The reactive distillation column 1 according to the invention is suitable for a reaction system in which at least one liquid-phase feed and at least one gas-phase feed are subjected to a chemical reaction over a hydrocracking catalyst and at least one gas-phase product is present in the reaction product. Such as hydrocracking of petroleum fractions and chemical synthesis oils, hydrodewaxing of diesel and lubricant oil fractions, hydrotreating of various petroleum fractions, and the like.
In the reactive distillation tower 1, each layer of tower tray comprises a downcomer, an overflow weir and a liquid receiving disc 16, a liquid sealing baffle plate is arranged on the tower tray, the liquid sealing baffle plate is connected with a gas phase channel, adjacent tower trays are separated by an inverted umbrella-shaped baffle plate, each layer of tower tray is of an annular structure, and the annular outer edge is connected with the gas phase channel. The gas phase channels are common channels for removing gas phase products generated by chemical reaction on each tray. The liquid feed locations of the embodiments of the present invention are all above the first tray, or may have liquid feed on some trays or on each tray, and gas phase feed has feed at the lower portion of each tray. The catalyst filling area is arranged above each layer of tray, the liquid phase feed radially flows through the catalyst bed, the gas phase feed enters from the lower part of the tray and reacts under the action of the catalyst, the gas phase material generated after the reaction directly breaks away from the reaction system and enters the gas phase channel at the outer side, and the liquid phase leaves the bed and enters the next bed through the downcomer. Because the reaction and the separation are carried out simultaneously, the reaction balance can be destroyed, and the conversion rate of reactants and the selectivity of target products can be effectively improved.
As shown in fig. 10, the raw oil 1001 is mixed with the recycle gas 1005 of the hydrotreatment stage and the additional fresh hydrogen 1006 by heat exchange, the mixture 1002 of the raw oil and the hydrogen enters the hydrocracking pretreatment reactor 102, a hydrocracking pretreatment catalyst is filled in the reactor, one or more beds can be provided, cold hydrogen can be injected between the beds, and the raw material undergoes the pre-reactions such as hydrodenitrogenation, hydrodesulphurisation, and hydro-saturation in the reactor 102. The pre-reaction effluent 1003 enters a high-pressure separator 104 for gas-liquid separation after heat exchange and cooling, the separated hydrogen-rich gas, namely a high-pressure separator gas-phase effluent 1004 of a hydrotreatment section, circulates to a hydrocracking pre-treatment reactor 102 through a circulating compressor 103, a liquid-phase effluent 1007 of the high-pressure separator 104 is heated by a liquid-phase feed heating furnace 105 of the hydrocracking reactor, is used as a feed oil 1008 of the hydrocracking reactor to be mixed with circulating gas 1014 (or can be used as a feed 1009 of the hydrocracking reactor after being mixed with the liquid-phase effluent 1007 of the high-pressure separator 104), enters a hydrocracking reactor 106 after being mixed with hydrogen of the circulating gas 1014, the hydrocracking reactor is used as a mixed feed 1009 of the hydrocracking reactor, the hydrocracking reactor 106 is filled with a hydrocracking catalyst, the number of beds is generally one, the materials are subjected to shallow hydrocracking reaction on the catalyst, the liquid-phase effluent 1007 enters the middle part of a catalytic reaction distillation tower 1, the upper section and the lower section are the upper section, the product is subjected to rectification separation, and the lower section is the hydrocracking reaction section. The top stream 1011 from the catalytic reaction distillation column passes through the heat exchange cooler 107 and the cooled mixture is fed to the reflux drum 108 as reflux drum feed 1012 for gas-liquid separation. The hydrogen-rich gas reflux tank top gas 1013 circulates in two paths, one path circulates to the hydrocracking reactor 106 through a circulating hydrogen pipeline 1014, the other path enters the circulating hydrogen heating furnace 109 through the catalytic distillation circulating hydrogen pipeline 1015 and is mixed with the supplementary hydrogen 1021, and enters the reactor through a catalytic reaction distillation tower hydrogen inlet pipeline 1016 (a plurality of pipelines, at least one catalyst bed layer) after heating. The liquid phase stream 1017 at the bottom of reflux drum 108 is divided into two parts, one part is returned to the top of the rectifying section of the catalytic distillation column 1 as reflux through a reflux line 1018 of the top of the catalytic distillation column, and the other part is taken as a light product of the device through a product line 1019 at the top of the catalytic distillation column. The lower section of the catalytic reaction distillation tower 1 is a reaction section and is a place for carrying out hydrocracking reaction on heavy fractions. Here, the light fraction produced by the hydrocracking reaction of the heavy fraction rises to the rectifying section through the gas phase passage, the uncracked heavy fraction passes through the catalyst bed in stages, and the heavy fraction still not cracked through the lowermost bed is discharged from the bottom product line 1020 of the catalytic reaction distillation column.
Examples 1 to 5
The process flow diagram in each of the following examples is the same as that in FIG. 10, the process conditions in each of the examples are shown in Table 2, and the properties of the obtained main products are shown in Table 3. In each example, the hydrotreating reactor is filled with FF-36 catalyst, the hydrocracking reactor is filled with FC-24 catalyst, the catalyst bed layer of the catalytic reaction unit of the catalytic reaction distillation tower is filled with FC-24 catalyst, and the raw oil uses catalytic diesel, and the properties are shown in Table 1. In each example, 1 catalyst bed is arranged in the hydrocracking pretreatment reactor. The hydrocracking reactor is internally provided with 1 catalyst bed layer. The number of catalyst beds in the catalytic reactive distillation column was 6, wherein the diameter to height ratio of each catalyst bed was 5:1 for example 1, 20:1 for example 2, 40:1 for example 3, 100:1 for example 4, and 500:1 for example 5. The inert filler in the catalytic reaction distillation tower is Raschig ring.
TABLE 1 oil Properties of raw materials
| Raw oil | Catalytic cracking diesel oil A | Catalytic cracking diesel oil B |
| Density/g.cm -3 | 0.9440 | 0.9213 |
| Distillation range/° C | 148~370 | 151~344 |
| Sulfur content, wt% | 1.10 | 1.05 |
| Nitrogen content/. Mu.g.g -1 | 910 | 860 |
| Aromatic hydrocarbon, wt% | 78.0 | 75.3 |
| Cetane number | 15.1 | 17.0 |
Comparative example 1
Adopts a conventional two-stage hydrocracking method, namely a refining and cracking method process. The cracking reactor adopts raw materials to enter from the upper part of the reactor, and hydrogen enters from the lower part of the reactor. The cracked distillate is separated by a separating tower and then sent out as a product. The aspect ratio of the catalyst beds in the cracking reactor was 1:3. The other conditions were the same as in example 1. The properties of the main products obtained are shown in Table 3.
Table 2 example process conditions
TABLE 3 Properties of the Main product
| Product(s) | Example 1 | Example 2 | Example 3 | Example 4 | Example 5 | Comparative example 1 |
| Naphtha (naphtha) | ||||||
| Yield, wt% | 51.3 | 52.6 | 53.2 | 56.3 | 60.1 | 50.6 |
| Distillation range, C | ||||||
| 50%/End point | 141/199 | 146/198 | 151/196 | 142/200 | 145/201 | 147/198 |
| Aromatic hydrocarbon content, wt% | ||||||
| RON | 94.5 | 94.3 | 94.6 | 93.4 | 93.3 | 89.1 |
| Diesel oil | ||||||
| Yield, wt% | 45.8 | 43.1 | 42.6 | 39.6 | 35.8 | 45.7 |
| Distillation range, C | ||||||
| Initial point/50% | 189/260 | 190/263 | 192/261 | 188/262 | 191/261 | 188/259 |
| Cetane number | 47.1 | 47.8 | 48.5 | 48.8 | 53.5 | 40.2 |
The above describes in detail the specific embodiments of the present invention, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.
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