CN113061466A

CN113061466A - Combined distillation method for generating oil by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking

Info

Publication number: CN113061466A
Application number: CN202110173227.6A
Authority: CN
Inventors: 何巨堂; 何艺帆
Original assignee: Luoyang Ruihua New Energy Technology Development Co ltd
Current assignee: Luoyang Ruihua New Energy Technology Development Co ltd
Priority date: 2021-02-04
Filing date: 2021-02-04
Publication date: 2021-07-02

Abstract

The combined distillation method of tail liquid circulating hydrocarbon material up-flow type hydrogenation thermal cracking generated oil is suitable for the working conditions of circulating reaction of reaction products containing wax oil components and residual oil components and at least part of heavy wax oil components, circulating reaction of most unconverted residual oil, and discharging of a small part of unconverted residual oil, such as residue boiling bed hydrocracking, residual oil suspension bed hydrocracking, direct coal slurry hydrogenation liquefaction and the like, and the operation temperature of the circulating reaction residual oil can be reduced as much as possible when the yield of the unconverted residual oil is reduced in the fractionation process; the unevaporated liquid material obtained by predistillation of hydrogenation produced oil is divided into two paths of SF1 and SF 2; the SF1 is subjected to shallow fractionation in the first vacuum distillation part to obtain a first residual oil containing a large amount of heavy wax oil components in the circulating reaction, so that the energy consumption can be reduced, and the thermal condensation can be inhibited; the SF2 is deeply evaporated in the second flash evaporation part to obtain the second residual oil with low concentration of wax oil components discharged outside and the second flash evaporation gas jointly recovered in the first vacuum distillation part, so that the process can be simplified, and part of the second residual oil can be recycled for reaction.

Description

Combined distillation method for generating oil by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking

Technical Field

The invention relates to a combined distillation method for generating oil by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking, which is suitable for the working conditions of circulating reaction of a reaction product containing a wax oil component and a residual oil component, circulating reaction of at least part of a heavy wax oil component, circulating reaction of most of unconverted residual oil, and discharging of a small part of unconverted residual oil, such as residue boiling bed hydrocracking, residual oil suspension bed hydrocracking, direct coal slurry hydrogenation liquefaction and the like, wherein the operation temperature of the circulating reaction residual oil can be reduced as much as possible when the yield of the unconverted residual oil is reduced in the fractionation process; the unevaporated liquid material obtained by predistillation of hydrogenation produced oil is divided into two paths of SF1 and SF 2; the SF1 is subjected to shallow fractionation in the first vacuum distillation part to obtain a first residual oil containing a large amount of heavy wax oil components in the circulating reaction, so that the energy consumption can be reduced, and the thermal condensation can be inhibited; the SF2 is deeply evaporated in the second flash evaporation part to obtain the second residual oil with low concentration of wax oil components discharged outside and the second flash evaporation gas jointly recovered in the first vacuum distillation part, so that the process can be simplified, and part of the second residual oil can be recycled for reaction.

Background

The invention relates to a vacuum distillation method for oil generated by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking reaction, which aims at the working conditions that the reaction product contains wax oil components and residual oil components, at least part of the heavy wax oil components are circularly reacted, most of unconverted residual oil is circularly reacted, and a small part of unconverted residual oil is discharged outside, such as residue boiling bed hydrocracking, residual oil suspension bed hydrocracking, direct coal slurry hydrogenation liquefaction and the like, and focuses on the main problems that: the method has the advantages of reducing the energy consumption in the vacuum fractionation process, simplifying the process, reducing the investment, reducing the amount of heavy wax oil in discharged residues, improving the yield of distilled oil, inhibiting coking, prolonging the operation period and reducing the severity of vacuum distillation operation, namely, the method realizes the seemingly contradictory problem of reducing the yield of heavy wax oil in unconverted residual oil in the vacuum fractionation process and simultaneously reducing the operation temperature in the process of separating the circulating reaction residual oil to the greatest extent.

For the heavy wax oil component, the conventional boiling point is between 480 and 550 ℃, so the distillation cut point is set between the medium wax oil and the heavy wax oil, and the distillation cut point is set between the heavy wax oil and the vacuum residue oil, the balance temperature difference in the vacuum evaporation process can reach 35 to 50 ℃, for example, the difference between the operation temperature of 325 ℃ and 365 ℃ is very obvious.

For the processes of hydrocracking of a vacuum residue boiling bed, hydrocracking of a vacuum residue suspension bed, direct liquefaction of coal slurry and the like, the asphaltene in the cracked produced oil belongs to a thermosensitive component and is easy to thermally condense or even coke at high temperature, the content of the asphaltene in the vacuum residue of the product of hydrocracking of the boiling bed is as high as 40-60 wt%, the content of the asphaltene in the vacuum residue of the product of hydrocracking of the vacuum residue suspension bed is as high as 30-40 wt% and contains a suspension bed hydrogenation catalyst for promoting thermal condensation, the content of the asphaltene in the vacuum residue of the product of direct liquefaction of coal slurry is as high as 50-60 wt% and contains a suspension bed hydrogenation catalyst for promoting thermal condensation, therefore, when the vacuum residue boiling bed hydrocracking of the vacuum residue and the vacuum residue hydrocracking of the vacuum residue are carried out by using a circulating reaction vacuum residue, the operating temperature and the high-temperature retention time of the fractionation process of the thermally cracked produced oil, the operating parameters which are directly related to the yield of the thermal condensate must be optimized as far as possible, i.e. the operating temperature is reduced and the residence time is shortened.

In fact, the task of the invention is to separate the same composition material into the same components (heavy wax oil component, residual oil component) and go to at least 2 different fractions, wherein most of the circulating reaction heavy wax oil component does not need to be separated from the circulating reaction residual oil, and a small part of the heavy wax oil component needs to be separated from discharged unconverted oil (the same component as the circulating reaction residual oil), namely classified vacuum distillation is needed, but the disadvantages of complicated flow and high investment, which require centralized fractionation of the same kind of materials, are avoided, and therefore, the principle flow of the invention is provided.

The idea of the invention is: the combined distillation method of tail liquid circulating hydrocarbon material up-flow type hydrogenation thermal cracking generated oil is suitable for the working conditions of circulating reaction of reaction products containing wax oil components and residual oil components and at least part of heavy wax oil components, circulating reaction of most unconverted residual oil, and discharging of a small part of unconverted residual oil, such as residue boiling bed hydrocracking, residual oil suspension bed hydrocracking, direct coal slurry hydrogenation liquefaction and the like, and the operation temperature of the circulating reaction residual oil can be reduced as much as possible when the yield of the unconverted residual oil is reduced in the fractionation process; the unevaporated liquid material obtained by predistillation of hydrogenation produced oil is divided into two paths of SF1 and SF 2; the SF1 is subjected to shallow fractionation in the first vacuum distillation part to obtain a first residual oil containing a large amount of heavy wax oil components in the circulating reaction, so that the energy consumption can be reduced, and the thermal condensation can be inhibited; the SF2 is deeply evaporated in the second flash evaporation part to obtain the second residual oil with low concentration of wax oil components discharged outside and the second flash evaporation gas jointly recovered in the first vacuum distillation part, so that the process can be simplified, and part of the second residual oil can be recycled for reaction.

In order to reduce the maintenance workload of the vacuum tower, the vacuum tower of the first vacuum distillation part can be decomposed into a feeding flash evaporator and a flash gas distillation tower (a rectifying tower), the flash gas distillation tower basically has no risk of coking, carbon deposition, material scouring and the like, the risks of coking, carbon deposition, material scouring and the like are concentrated in the feeding flash evaporator, and the maintenance and centralized management are facilitated.

The invention forms the operation modes of classified evaporation and concentrated rectification of flash steam, and combines shallow reduced pressure distillation and deep reduced pressure distillation.

In order to simplify the process and reduce gas transfer pipelines, a set of equipment of a dividing wall vacuum distillation tower can be used for completing the task of the invention, and the invention is a feasible selection for the working condition that the coking of residual oil components is not serious.

Regarding the vacuum distillation method of oil generated by the upflow hydrogenation thermal cracking reaction of hydrocarbon material, the existing several methods have obvious defects:

scheme-single shallow vacuum distillation scheme

In order to prevent the bottom temperature of the vacuum distillation tower from being too high, which causes serious thermal condensation of the bottom oil of the vacuum distillation tower and deteriorates the quality of the bottom oil of the vacuum distillation tower in a circulating reaction, the bottom temperature of the vacuum distillation tower is artificially reduced to cause the bottom oil of the vacuum distillation tower to contain more heavy wax oil, so that the discharged unconverted oil UCO also contains more heavy wax oil, the distillation rate of the wax oil in the distillation process is reduced, the yield of the discharged unconverted oil UCO is increased, the value of the discharged unconverted oil UCO used as asphalt or as a coal substitute is greatly reduced, huge economic loss is formed, the price difference between the wax oil and asphalt is about 1500-2000 yuan/ton, a 200 million tons/year residual oil hydrocracking device is adopted, the yield of the discharged unconverted oil UCO is increased from 8 weight percent to 10 weight percent and differs by 4 million tons/year, and the yield of the wax oil is reduced by 4 million tons/year, The yield of the asphalt is increased by 4 ten thousand tons per year, the price difference is 6000-8000 ten thousand yuan per year, the loss is serious, and the scheme is not advisable;

② scheme II, namely a single-depth reduced pressure distillation scheme

In order to prevent the discharged unconverted oil UCO from containing more heavy wax oil to form huge economic loss, the single deep reduced pressure distillation scheme is adopted, and the defects are that: firstly, the outlet temperature of a feeding heating furnace of a decompression tower is overhigh, and fuel is wasted (heavy wax oil which causes a large amount of cyclic reaction is repeatedly evaporated and condensed); secondly, thermal condensation of the bottom oil of the vacuum distillation tower and the feeding heating furnace tube of the vacuum tower is intensified, so that the bottom oil of the vacuum distillation tower subjected to the circulating reaction is gradually deteriorated in quality and seriously scaled or coked, the continuous operation period of the feeding heating furnace of the vacuum tower and the vacuum tower is shortened, the operating rate of the device is reduced, and the overhaul workload of workers is increased; thirdly, the content of asphaltene in a liquid phase of R10 in the upflow type hydrogenation thermal cracking reaction process of the hydrocarbon material is increased, the reaction effect is deteriorated, the overall conversion rate of residual oil is reduced, and the economic benefit is greatly reduced; therefore, the single deep reduced pressure distillation scheme cannot realize the process target of pure deep drawing of the wax oil without causing other risks; in fact, the short-term deep drawing target can be realized only at the cost of scaling and coking, and the operation risk is huge; under the threat of operational risk, the single deep vacuum distillation scheme would fall back to the single shallow vacuum distillation scheme;

(III) patent of the third patent of Zl201080035430.7

Chinese patent ZL201080035430.7 is directed to a method and apparatus for separating bitumen from slurry hydrocracked vacuum gas oil and compositions by slurrying a heavy hydrocarbon feed with a particulate solid material to form a heavy hydrocarbon slurry and hydrocracking in a slurry hydrocracking unit to produce Vacuum Gas Oil (VGO) and bitumen. The first vacuum column separates VGO from pitch and the second vacuum column further separates VGO from pitch. As much as 15 wt-% of the VGO can be recovered by the second vacuum column and recycled to the slurry hydrocracking unit while yielding a pitch composition that can be pelletized and transported without sticking together;

in chinese patent ZL201080035430.7, the reaction product of slurry hydrocracking reactor 20 is separated into hot high-pressure oil 34 in hot high-pressure separator 30; the hot high-fraction oil 34 enters a hot flash tank 36 to separate out hot low-fraction oil; the hot low fraction oil enters a fractionating section 50 to separate out bottom oil of the prefractionation tower; the prefractionating tower bottom oil is heated by a heating furnace 84 and then enters a first vacuum tower 90 to separate out first vacuum tower bottom oil containing heavy wax oil and residual oil; the first vacuum column bottoms, with or without heating by pre-evaporator 102, enters large diameter flash section 108 of second vacuum column 100 (evaporator stripper 100), vapor VGO may exit the top of the evaporator stripper through an entrainment separator such as a demister to separate condensables, vapor exits in line 110 and enters condenser 112 and, where applicable, accumulator 114, where vacuum may be drawn from condenser 112; VGO, mainly heavy wax oil component HVGO, withdrawn in line 116 is recycled to the SHC reactor 20 via line 8; the SHC reactor 20 is in downstream communication with the top of the second vacuum column 100; a portion of the HVGO in line 116 can be recovered and flowed out as net product in line 124; the pitch in the second vacuum column 100 cascades down on heated or unheated trays, such as tube tray trays, while the remaining volatiles are stripped by the rising vapor, the trays providing a new liquid film at each stage, renewing the surface of the pitch film for evaporation and stripping; the trays may define an internal cavity in communication with a heating fluid from pipe 126 for indirectly heating the pitch passing through the trays; the heated fluid exits the second vacuum tower 100 in line 128 for reheating and recycling; an inert gas such as steam or nitrogen may be distributed into the column from line 118 to strip the pitch and further enhance mass transfer, which is a membrane-generating vaporizer; the second pitch stream exits the second vacuum column 100 in line 120; the second pitch stream comprises less than 14 wt.% VGO, typically no more than 13 wt.% VGO, preferably no more than 10 wt.% VGO, but typically the second pitch stream comprises at least 1 wt.% VGO, which is in a hydrocarbon fraction having a conventional boiling range of 300 ℃ to 538 ℃;

in contrast to the conventional shallow vacuum distillation scheme, chinese patent ZL201080035430.7, provides a second vacuum column 100 operating in series downstream of the first vacuum column 90, with the aim of recovering as much as 15 wt.% of VGO (mainly heavy VGO) from the first vacuum column bottoms pitch for use as product or recycle reaction heavy VGO; it can be seen that the second vacuum tower 100 operated in series is provided, not only the investment is increased, but also the cyclic reaction heavy VGO is completely subjected to the evaporation process of the second vacuum tower 100 and is used as distillate oil to perform reaction cycle, rather than the majority of the VGO is subjected to the cyclic reaction along with the cyclic reaction residual oil without the evaporation process, and the waste of heat energy is serious; it can also be seen that the asphalt at the bottom of the first vacuum tower, after at least 1 heating process (in the film formation evaporator), even 2 heating processes (in the pre-evaporator 102 and the film formation evaporator respectively), is heated for several times, which may seriously aggravate the thermal condensation reaction of the residual oil in the bottom oil of the vacuum tower, and is reasonable and acceptable as the process target of chinese patent ZL201080035430.7, namely, the asphalt composition which is made into particles and is not stuck together during transportation, but is unacceptable for the process target of the present invention for carrying out the circulating reaction of the unconverted residual oil; chinese patent ZL201080035430.7 realizes the aim of deep drawing heavy wax oil at the cost of deteriorated final asphalt quality, and can not realize large-amount circulation of unconverted residual oil, so that the technical scheme of the invention has no comparability;

the technical scheme of the invention is different from the Chinese patent ZL201080035430.7 in that:

firstly, in the chinese patent ZL201080035430.7, the prefractionation column bottom oil separated from the fractionating section 50 adopts a method of 2-stage series distillation (at least one heating process exists before or during each stage of distillation) and respectively recovering flash steam, and the residual oil component is heated for many times and has long heating time;

the method for producing the bottom oil of the oil pre-distillation tower by hydrogenation adopts a method of branching parallel flash evaporation (the most one heating is probably needed before the flash evaporation) and combined rectification of flash evaporation steam, so that the residual oil component is heated for a few times, the heating time is short, and only one fractionating tower is needed;

secondly, the unevaporated oil in the vacuum distillation process of the Chinese patent ZL201080035430.7 is totally taken as asphalt to remove solid and form, and is not (and cannot be) used for circulating reaction;

most of the unevaporated oil in the reduced pressure distillation process is used as the tail oil of the circular reaction, and a small part of the unevaporated oil is used as the discharged unconverted oil, so that the use value is high;

thirdly, in the Chinese patent ZL201080035430.7, at least 15 wt% of heavy wax oil component in the tower bottom oil of the first vacuum tower 90 is evaporated and separated in the second vacuum tower 100 to be used as distillate oil heavy wax oil, and then the circulation reaction can be carried out; in the vacuum distillation process, most of the heavy wax oil components do not need to be subjected to an evaporation process, and are directly subjected to a circulating reaction along with the circulating reaction residual oil components, so that the fractionation load is reduced, and the heat energy can be saved;

(scheme IV) Chinese patent ZL201010217358.1

The coal tar suspension bed hydrogenation method of the heterogeneous catalyst ZL201010217358.1 comprises the steps of coal tar raw material pretreatment, distillation and separation, coal tar heavy fraction suspension bed hydrocracking and light distillate oil conventional upgrading process, wherein the suspension bed hydrogenation reaction temperature is 320-480 ℃, the reaction pressure is 8-19 MPa, and the volume space velocity is 0.3-3.0 h^-1The volume ratio of hydrogen to oil is 500-2000, the catalyst is a powdery granular coal tar suspension bed hydrogenation catalyst containing a single metal active component or a composite multi-metal active component of molybdenum, nickel, cobalt, tungsten or iron, the addition amount is that the weight ratio of the metal amount of the active component to the coal tar raw material is 0.1: 100-4: 100, most of tail oil containing the catalyst after light oil is separated from a hydrogenation reaction product is directly circulated to a suspension bed reactor, and a small part of tail oil is subjected to catalyst removal treatment and then is circulated to the suspension bed reactor, so that the aim of producing light oil at the maximum amount from the coal tar and recycling the catalyst is fulfilled, and the utilization efficiency of the raw material and the catalyst is greatly improved;

in the chinese patent ZL201010217358.1, after the reaction effluent of the suspension bed hydrogenation reactor is separated by a high temperature separator and a low temperature separator, a liquid-solid phase high-low oil mixture stream and a hydrogen-rich gas are obtained, the hydrogen-rich gas is used as recycle hydrogen, the liquid-solid phase high-low oil mixture flows through a normal pressure fractionating tower, light distillate oil with a temperature of less than 370 ℃ is obtained at the tower top, heavy oil with a temperature of more than 370 ℃ is obtained at the tower bottom, most of the heavy oil with a temperature of more than 370 ℃ is directly circulated into the suspension bed hydrogenation reactor as recycle oil for further hydrogenation and lightening reaction, the circulation amount is two thirds to four fifths of the heavy oil with a temperature of the normal bottom, the rest part of the heavy oil with a temperature of more than 370 ℃ enters a solid-liquid separation system for solid-liquid separation, the solid-liquid separation can be in the form of filtration or centrifugal separation or vacuum fractionation, the catalyst residue and the suspension bed hydrogenation heavy, circularly entering a suspension bed hydrogenation reactor for further hydrogenation and lightening reaction; throwing or regenerating the separated catalyst;

according to the description of the Chinese patent ZL201010217358.1, a small part of normal bottom heavy oil enters a solid-liquid separation system, when the solid-liquid separation system adopts a vacuum fractionation mode, catalyst residues (such as vacuum tower bottom oil) and suspension bed hydrogenation heavy distillate oil (such as vacuum tower distillate oil) are obtained after separation, the heavy distillate oil (such as vacuum tower distillate oil) circularly enters a suspension bed hydrogenation reactor to further carry out hydrogenation and lightening reaction, and the suspension bed hydrogenation heavy distillate oil is not described to be sent out as a product; the removed catalyst (such as bottom oil of a vacuum tower) is thrown (discharged) or regenerated;

in chinese patent ZL201010217358.1, because the liquid-solid phase high-low fraction oil mixture is fractionated in a normal pressure fractionating tower, the normal bottom heavy oil containing a catalyst obtained at the bottom of the tower and having a temperature higher than 370 ℃ inevitably contains a large amount of light wax oil, medium wax oil and heavy wax oil, and therefore, two-thirds to four-fifths of the normal bottom heavy oil is used as a circulating reaction material for circulating reaction, which means that most of the light wax oil, medium wax oil and heavy wax oil are all subjected to circulating hydrocracking, and the light wax oil and the medium wax oil are not the object of the present invention, because the light wax oil and the medium wax oil are converted by a suspension bed or boiling bed hydrocracking manner, the scale of the apparatus is increased, and at the same time, the gas yield is too high, the hydrogen consumption is too high, and the conversion product still needs to be subjected to fixed bed hydrogenation upgrading, so that the investment and the energy consumption are greatly increased, and the method is unreasonable; in the Chinese patent ZL201010217358.1, the light wax oil product and the medium wax oil product are directly introduced into a fixed bed hydrocracking device (containing a hydrofining section) for treatment, and the point is not comparable with the technical scheme of the invention;

in the chinese patent ZL201010217358.1, the rest of the extra heavy oil enters the solid-liquid separation system for solid-liquid separation, and the solid-liquid separation can be performed in the form of filtration, centrifugal separation or vacuum fractionation, but the technical details of the vacuum fractionation process of the rest of the extra heavy oil aiming at solid-liquid separation are not described, and the similar scheme of the present invention is not disclosed, and the object of the present invention cannot be achieved.

The method of the invention is not reported.

The invention aims to provide a combined distillation method for producing oil by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking

Disclosure of Invention

The invention discloses a combined distillation method for generating oil by tail liquid circulating hydrocarbon upflow type hydrogenation thermal cracking, which is characterized by comprising the following steps:

(1) in the upflow type hydrogenation thermal cracking reaction process R10, under the conditions that hydrogen and conventional liquid hydrocarbon exist or do not exist at the same time, and a mixed phase material of solid particles exists or does not exist, hydrogenation reaction R10-R is carried out on the hydrocarbon material R10F at least containing carbon element and hydrogen element, at least part of hydrogenation thermal cracking reaction is carried out, and the hydrocarbon material is converted into a hydrogenation reaction product R10P;

a hydrocarbon feed R10F comprising a liquid feedstock R10FL, with or without solid particulate hydrocarbon feed R10 FS;

liquid raw material R10FL, comprising fresh liquid raw material R10FL-NEW, total circulating reaction tail liquid U20-RR from step (4);

the hydrogenation reaction product R10P comprises a hydrocarbon component with a conventional boiling point lower than 480 ℃, a heavy wax oil component with a conventional boiling point between 480 and 550 ℃, and a residual oil component with a conventional boiling point higher than 550 ℃, wherein the residual oil component with a conventional boiling point higher than 550 ℃ is a hydrogenation converted product of the hydrocarbon feed R10F and/or a component carried by the hydrocarbon feed R10F;

hydrogenation reactions R10-R, hydrocracking reactions comprising at least a portion of the liquid feedstock R10FL, hydroliquefaction reactions with or without solid particulate hydrocarbon feedstock R10 FS;

in the process of hydrogenation reaction R10-R, catalyst R10-CAT exists or does not exist, catalyst R10-CAT contains or does not contain molybdenum sulfide particles, and catalyst R10-CAT contains or does not contain Fe7S 8;

in the upflow type hydrogenation thermal cracking reaction process R10, hydrogen-donating hydrocarbon or a hydrogen-donating hydrocarbon precursor containing a hydrocarbon component with a conventional boiling point of 230-400 ℃ is present or not, an external hydrogen-donating hydrocarbon material stream OUT-TO-R10-DS containing hydrogen-donating hydrocarbon SH with a conventional boiling point of 230-400 ℃ is used or not is used, and a hydrogen-donating hydrocarbon component R10PRO-DSC with a conventional boiling point of 230-400 ℃ is generated or not is generated;

a hydrogenation reactor R10-XE is used in the upflow type hydrogenation thermal cracking reaction process R10 of the hydrocarbon material, a hydrogenation reaction product R10P is a mixed-phase material containing hydrogen and conventional liquid hydrocarbon and containing or not containing solid particles, the mixed-phase material is in the form of 1-path or 2-path or multi-path material R10P-X when leaving the reactor R10-XE, and the material R10P-X is a gas phase or liquid phase or gas-liquid mixed-phase or gas-liquid-solid three-phase material flow;

a hydrocarbon material up-flow type hydrogenation thermal cracking reaction process R10, at least one up-flow type hydrogenation thermal cracking reactor is used;

(2) in the thermal high-pressure separation process, separating a hydrogenation reaction product R10P to obtain a wax-containing oil component in a high-temperature state, thermal high-molecular oil HHPL containing a residual oil component and thermal high-molecular gas HHPV mainly composed of hydrogen in volume;

a material containing a wax oil component and a residual oil component based on hot high-molecular-weight oil HHPL, which is used as a feed U10-F in the pre-distillation process U10 in the step (3);

when a hot low-pressure separation process is set, the hot high-separation oil HHPL is separated into hot low-separation oil and hot low-separation gas in the hot low-pressure separation process; a hot low oil based waxy oil component, resid component containing feed, U10-F, as feed U10 for the pre-distillation process of step (3);

(3) in pre-distillation process U10, pre-distillation process U10 is fed U10-F, separated into gas U10-G, a pre-distillation non-evaporated oil U10-D containing a waxy oil component and a resid fraction, and a pre-distillation narrow distillate oil U10-LP, with or without the presence, using a step comprising a pre-distillation column U10T;

a stream of unevaporated oil U10-D based on the pre-distillation process, U20-F as feed to combined vacuum distillation process U20;

(4) in the combined reduced pressure distillation process U20, carrying out fractionation of a combined reduced pressure distillation process U20 feeding U20-F to obtain discharged unconverted oil UCO, total circulating reaction tail liquid U20-RR and other narrow fraction distillate oil;

a combined vacuum distillation process U20 comprising a first vacuum distillation section U21, a second flash section U22;

a wax-containing oil component, a resid-containing component based on combined vacuum distillation process U20 feed U20-F, divided into a first feed SF1 and a second feed SF 2;

firstly, in a first reduced pressure distillation part U21, a first feed SF1 is subjected to flash evaporation separation in a first flash evaporation process to obtain a first feed flash evaporation gas SF1V and a first feed flash evaporation liquid SF1L containing wax oil components and dreg oil components, and a liquid phase flashed out of the first feed SF1 is subjected to gas stripping with or without using a gas stripping gas, wherein the gas stripping purpose is to obtain a first feed flash evaporation liquid SF1L after stripping out part of low-boiling-point components;

materials based on the first feeding flash liquid SF1L are used as a circulating reaction tail liquid SF1L-RR and return to the upflow hydrocracking reaction process R10 for circulating reaction;

separating 1-path or 2-path or multi-path narrow fraction oil SF1V-LP from the material based on the first feeding flash gas SF1V through the rectification section of the first vacuum distillation tower; the narrow-cut oil SF1V-LP contains a low-boiling-point wax oil component with a conventional boiling point of 350-480 ℃;

secondly, in a second flash evaporation part U22, the second feed SF2 is flash-evaporated and separated into second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF2L containing residual oil components in a second flash evaporation process, and gas stripping is carried out on a liquid phase flash evaporated from the second feed SF2 with or without using gas stripping gas, wherein the gas stripping aims at obtaining second feed flash evaporation liquid SF2L after stripping off part of low-boiling-point components;

the weight concentration of the fraction having a conventional boiling point below 550 ℃ in second feed flash SF2L is lower than the weight concentration of the fraction having a conventional boiling point below 550 ℃ in first feed flash SF 1L;

at least a portion of the second feed flash SF2L, which is used as external unconverted oil UCO, with or without a stream based on the second feed flash SF2L, being returned as recycled reaction tail SF2L-RR to the upflow hydrocracking reaction process R10 for recycle reaction;

separating and recovering materials based on the second feeding flash gas SF2V through at least one part of the rectifying section of the first reduced pressure distillation tower; mixing a material based on the second feeding flash gas SF2V and a material based on the first feeding flash gas SF1V in a rectification section of the first vacuum distillation tower;

the total circulating reaction tail liquid U20-RR comprises circulating reaction tail liquid SF1L-RR and circulating reaction tail liquid SF 2L-RR.

In the present invention, the first flash process may be performed in a first reduced pressure distillation column:

(4) in the combined reduced pressure distillation process U20, firstly, in a first reduced pressure distillation part U21, a first feeding SF1 enters a first reduced pressure distillation tower to carry out a first flash distillation process, and flash separation is carried out to obtain a first feeding flash evaporation gas SF1V and a first feeding flash evaporation liquid SF 1L;

the first feed flash gas SF1V is separated into narrow distillate oil SF1V-LP in the rectifying section of the first vacuum distillation tower.

According to the present invention, the first flash process may be performed in a first flash vessel outside the first reduced pressure distillation column:

(4) in the combined vacuum distillation process U20:

first, in a first vacuum distillation part U21, a first feed SF1 is subjected to a first flash process in a first flash vessel and separated into a first feed flash gas SF1V and a first feed flash liquid SF 1L;

the first feeding flash gas SF1V enters a first vacuum distillation tower to be separated into narrow-cut oil products SF 1V-LP.

According to the invention, (4) the operating conditions of the combined vacuum distillation process U20 can be selected from 1 of the following:

selecting one, namely, in a first reduced pressure distillation part U21, the concentration of hydrocarbon components with the conventional boiling point of 350-550 ℃ in a first feeding flash liquid SF1L is more than or equal to 17 wt%; ② in the second flash part U22, the concentration of the components with the conventional boiling point lower than 550 ℃ in the second feed flash liquid SF2L is less than or equal to 7 percent by weight;

selecting two, namely, in a first reduced pressure distillation part U21, the concentration of hydrocarbon components with the conventional boiling point of 350-550 ℃ in a first feeding flash liquid SF1L is more than or equal to 25 wt%; ② in the second flash part U22, the concentration of the components with the conventional boiling point lower than 550 ℃ in the second feed flash liquid SF2L is less than or equal to 2 weight percent;

selecting three, namely, in the first reduced pressure distillation part U21, the temperature of the first feeding flash liquid SF1L is T1; ② in the second flash section U22 the temperature of the second feed flash SF2L is T2, T2 is at least 20 ℃ above T1;

selecting four, namely, in the first reduced pressure distillation part U21, the temperature of the first feeding flash liquid SF1L is T1; ② in the second flash section U22 the temperature of the second feed flash SF2L is T2, T2 is at least 35 ℃ above T1;

fifthly, in a first reduced pressure distillation part U21, the first feeding SF1 enters a first flash process without being heated by a heating furnace to be separated into first feeding flash evaporation gas SF1V and first feeding flash evaporation liquid SF1L through flash evaporation; in a second flash evaporation part U22, after the temperature of second feed flash evaporation liquid SF2L is raised by a second heating furnace, the second feed flash evaporation liquid enters a second flash evaporation process and is subjected to flash evaporation separation to obtain second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF 2L;

selecting sixth, in the first reduced pressure distillation part U21, after the temperature of the first feeding SF1 is raised by a first heating furnace, the first feeding SF1 enters a first flash evaporation process to be separated into a first feeding flash evaporation gas SF1V and a first feeding flash evaporation liquid SF1L by flash evaporation, and the temperature of the first feeding SF1 discharged out of a furnace tube of the first heating furnace is FT 1; in a second flash evaporation part U22, after the temperature of second feed SF2 is raised by a second heating furnace, the second feed SF2 enters a second flash evaporation process to be subjected to flash evaporation separation to obtain second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF2L, the temperature of the second feed SF2 discharged out of a furnace tube of the second heating furnace is FT2, and the temperature FT2 is at least 20 ℃ higher than the temperature FT 1;

seventhly, in the first reduced pressure distillation part U21, after the temperature of the first feeding SF1 is raised by a first heating furnace, the first feeding SF1 enters a first flash evaporation process to be subjected to flash evaporation separation to obtain first feeding flash evaporation gas SF1V and first feeding flash evaporation liquid SF1L, and the temperature of the first feeding SF1 discharged out of a furnace tube of the first heating furnace is FT 1; in a second flash evaporation part U22, after the temperature of second feed SF2 is raised by a second heating furnace, the second feed SF2 enters a second flash evaporation process to be subjected to flash evaporation separation to obtain second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF2L, the temperature of the second feed SF2 discharged out of a furnace tube of the second heating furnace is FT2, and the temperature FT2 is at least 35 ℃ higher than the temperature FT 1;

eighthly, in the first reduced pressure distillation part U21, the liquid phase flashed out from the first feeding SF2 is not subjected to gas stripping; ② in the second flash section U22, the liquid phase flashed off from the second feed SF2 was stripped using a stripping gas.

In the present invention, the operation mode of the combined vacuum distillation process U20 can be as follows:

(4) in the combined reduced pressure distillation process U20, firstly, in a first reduced pressure distillation part U21, a first feed SF1 is heated by a first heating furnace and enters a first flash evaporation process to be subjected to flash evaporation separation to obtain a first feed flash evaporation gas SF1V and a first feed flash evaporation liquid SF 1L; in a second flash evaporation part U22, after the temperature of second feed SF2 is raised by a second heating furnace, the second feed SF2 enters a second flash evaporation process to be subjected to flash evaporation separation to obtain second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF 2L;

the first heating furnace and the second heating furnace are arranged as a combined heating furnace, the hearth of the first heating furnace and the hearth of the second heating furnace are separated by a partition wall, and the flue gas of the first heating furnace and the flue gas of the second heating furnace are mixed and then subjected to flue gas heat recovery.

The invention, (4) combining feed U20-F of the vacuum distillation process U20 into first feed SF1 and second feed SF2, the ratio of the weight flow rate SF1-WR of the first feed SF1 to the weight flow rate SF1-WR of the second feed SF2 being K10;

k10 (SF1-WR)/(SF2-WR), the value of K10 may be selected from one of the following:

①(97∶3)～(93∶7)；

②(93∶7)～(88∶12)；

③(88∶12)～(76∶24)。

selecting one, namely, in a first reduced pressure distillation part U21, the operating pressure of a first flash process is 0.045-0.033 MPa (absolute pressure), the tower top pressure of a first reduced pressure distillation tower is 0.015-0.003 MPa (absolute pressure), the temperature of a first feeding flash liquid SF1L is T1, and the temperature of T1 is 300-340 ℃; secondly, in a second flash evaporation part U22, the operating pressure of the second flash evaporation process is 0.045-0.033 MPa (absolute pressure), the temperature of second feeding flash evaporation liquid SF2L is T2, and the difference of the temperature T2 higher than the temperature T1 is 20-60 ℃;

and secondly, fractionating the first feed flash gas SF1V and the second feed flash gas SF2V in a first vacuum distillation tower to obtain a heavy wax oil fraction containing a fraction with a conventional boiling point of 480-550 ℃ and a low-boiling-point wax oil fraction containing a fraction with a conventional boiling point of 350-480 ℃.

In the present invention, (4) in the combined vacuum distillation process U20, a combined vacuum distillation of the first feed SF1 and the second feed SF2 may be performed using one combined vacuum column;

a flash evaporation zone is arranged at the lower part of the combined decompression tower and is separated into a first flash evaporation zone and a second flash evaporation zone by a partition plate, the first feeding material SF1 enters the first flash evaporation zone and is separated into a first feeding flash evaporation gas SF1V and a first feeding flash evaporation liquid SF1L, and the second feeding material SF2 enters the second flash evaporation zone and is separated into a second feeding flash evaporation gas SF2V and a second feeding flash evaporation liquid SF2L in a flash evaporation manner;

separating the first feed flash gas SF1V and the second feed flash gas SF2V into narrow fractions in a rectification section of the combined vacuum tower; mixing a material based on the first feed flash gas SF1V and a material based on the second feed flash gas SF2V in a rectification section of the combined vacuum column;

the first feed flash liquid SF1L is discharged from the first flash zone and is discharged from the combined decompression tower;

the second feed flash SF2L exits the second flash zone and exits the combined vacuum column.

In the invention, (4) in the combined vacuum distillation process U20, the working mode of arranging a flash evaporation zone at the lower part of the combined vacuum tower can be selected from 1 of the following:

alternatively, the first feed flash SF1L and the second feed flash SF2L are not mixed in the combined vacuum column;

selecting part of second feeding flash liquid SF2L as discharged unconverted oil UCO, discharging the second flash area and the combined decompression tower; in the combined decompression tower, part of the second feed flash liquid SF2L is discharged into a first flash zone as second feed flash liquid SF2L-RL for cyclic reaction, and is discharged out of the first flash zone after being mixed with the first feed flash liquid SF 1L;

the second feed flash liquid SF2L-RL for the cyclic reaction is discharged into the first flash zone in a manner that the elevation of the discharge port of the second feed flash liquid SF2L-RL for the cyclic reaction discharged from the second flash zone is higher than the operating elevation of the gas-liquid interface of the liquid phase aggregation zone of the first flash zone, for example, by at least 1.5 meters, so that the first feed flash liquid SF1L is prevented from entering the second flash zone and being mixed with the second feed flash liquid SF 2L-RL; alternatively, the recycle reaction is discharged into the first flash zone with second feed flash SF2L-RL at a level higher than the level of the discharge of the second flash zone with second feed flash SF2L-RL at the feed inlet of the first feed SF1 into the first flash zone, for example at least 1.5 meters higher, to prevent the liquid phase in the first feed SF1 from entering the second flash zone and mixing with the second feed flash SF 2L-RL.

According to the invention, (1) the hydrocarbon material R10F can be selected from 1 of the following materials:

liquid raw material R10FL mainly comprises hydrocarbons with normal boiling point higher than 450 ℃, and at least comprises a part of hydrocarbon components with normal boiling point higher than 550 ℃;

② the liquid raw material R10FL mainly comprises hydrocarbons with the conventional boiling point higher than 550 ℃;

③ the solid particulate hydrocarbon material R10FS comprises coal fines.

In the present invention, the operating conditions may be:

(1) the method comprises the following steps of (1) carrying out an upflow type hydrogenation thermal cracking reaction process R10 on a hydrocarbon material, wherein the reaction pressure is 6.0-30.0 MPa, and the reaction temperature is 350-480 ℃;

(2) in the hot high-pressure separation process, the operating pressure is 6.0-30.0 MPa, and the operating temperature is 350-450 ℃;

(3) the pre-distillation process U10 uses a pre-distillation column U10T operating pressure of 0.12-0.80 MPa (absolute pressure).

In the invention, (1) the hydrocarbon material up-flow type hydrocracking reaction process R10 can be a vacuum residue hydrocracking reaction process, and the weight hydrocracking rate of the components with the conventional boiling point higher than 550 ℃ in the liquid raw material R10FL can be selected from one of the following:

first, 90-97 wt%;

(ii) 80 to 90 wt%;

(iii) 70-80 wt%;

and 50-70 wt%.

The invention, (1) the hydrocarbon material up-flow type hydrogenation thermal cracking reaction process R10 can be a vacuum residue hydrogenation thermal cracking reaction process, the liquid raw material R10FL comprises a fresh vacuum residue raw material R10FL-NEW and total circulating reaction tail liquid U20-RR from the step (4);

the ratio of the weight flow rate of the hydrocarbon component with the conventional boiling point higher than 550 ℃ in the total circulating reaction tail liquid U20-RR to the weight flow rate R10FL-NEW-WR of the fresh vacuum residue raw material R10FL-NEW is K50;

k50 ═ U20-RR-VD-WR)/(R10FL-NEW-WR), K50 values were selected from one of the following:

①1.0～2.0；

②0.5～1.0；

③0.30～0.5；

the ratio of the weight flow rate of the heavy wax oil component with the conventional boiling point between 480-550 ℃ in the total circulating reaction tail liquid U20-RR to the weight flow rate R10FL-NEW-WR of the fresh vacuum residue raw material R10FL-NEW in the total circulating reaction tail liquid U20-RR is K70;

k70 ═ U20-RR-HVGO-WR)/(R10FL-NEW-WR), K70 values were selected from one of the following:

①0.5～0.8；

②0.3～0.5；

③0.15～0.3。

in the invention, (1) the working mode of the hydrocarbon material hydrocracking reactor used in the hydrocarbon material upflow hydrocracking reaction process R10 can be selected from 1 or more of the following modes:

firstly, a suspension bed reactor is a slurry bed reactor;

② a fluidized bed reactor;

③ a combined reactor of a suspension bed and a fluidized bed;

fourthly, micro-expansion bed.

According to the invention, (4) in the combined reduced pressure distillation process U20, the second flash evaporation part U22, the second feed SF2 in the second flash evaporator is subjected to the second flash evaporation process, and flash evaporation separation is carried out to obtain second feed flash evaporation gas SF2V and second feed flash evaporation liquid SF2L containing the residual oil component;

a heavy wax oil removing condensation section U20KD can be arranged in the second flash evaporator, the second feeding flash evaporation gas SF2V goes upward to pass through the heavy wax oil removing condensation section U20KD and is in countercurrent contact with the wax-containing oil cooling liquid U20QL entering from the upper part of the heavy wax oil removing condensation section U20KD for temperature reduction, at least part of heavy wax oil components in the second feeding flash evaporation gas SF2V are condensed into heavy wax oil removing liquid U20-HVGO which is discharged out of the second flash evaporator and is not mixed with the second feeding flash evaporation liquid SF 2L;

at least one part of the heavy component removal liquid U20-HVGO returns or does not return to the circulation reaction of the upflow hydrocracking reaction process R10;

the wax-containing oil cooling liquid U20QL is or is not a cooled and cooled heavy component removal liquid U20-HVGO and is or is not wax-containing oil liquid material discharged from the rectifying section of the first vacuum distillation tower of the first vacuum distillation part U21;

the gas leaving the heavy wax oil removing condensation section U20KD is taken as heavy gas and enters the rectification section of the first vacuum distillation tower of the first vacuum distillation part U21 for recovery; this can reduce the energy consumption for separating the heavy wax oil component from the low boiling point wax oil component in the second feed flash gas SF2, and prevent the first feed flash gas SF1V from being contaminated with the vacuum residue component, because the second feed flash gas SF2 generally contains a small amount of vacuum residue component.

Detailed Description

The present invention is described in detail below.

The pressure in the present invention refers to absolute pressure.

The conventional boiling point of the invention refers to the vapor-liquid equilibrium temperature of a substance at one atmospheric pressure.

The conventional boiling range as referred to herein refers to the conventional boiling range of the distillate fraction.

The compositions or concentrations or amounts or yield values of the components described herein are weight basis values unless otherwise specified.

In the upflow hydrogenation reactor, the macroscopic flow leading direction of the process medium in the reaction space or the hydrogenation catalyst bed layer is from top to bottom.

The upflow type expanded bed reactor is a vertical upflow type reactor, and belongs to an expanded bed reactor when a catalyst is used; the vertical type means that the central axis of the reactor is vertical to the ground in a working state after installation; the upflow means that the material main body flows in the reaction process from bottom to top to pass through the reaction space or the catalyst bed layer or flow in the same direction with the upward catalyst; the expanded bed means that a catalyst bed layer is in an expanded state in a working state, the expansion ratio of the catalyst bed layer is defined as the ratio KBED of the maximum height CWH of the working state when a reaction material passes through the catalyst bed layer and the height CUH of an empty bed standing state of the catalyst bed layer, generally, when the KBED is lower than 1.10, the bed is called a micro-expanded bed, when the KBED is between 1.25 and 1.55, the bed is called an ebullated bed, and a suspended bed is considered as the most extreme form of the expanded bed.

The back-mixed flow expanded bed reactor refers to an operation mode of using a reaction zone or a main reaction zone of the expanded bed reactor, wherein liquid flow back-mixing or circulating liquid exists; the return flow or the circulating liquid refers to at least one part of liquid phase XK-L in the intermediate product XK or the final product XK at the flow point K as a circulating liquid flow XK-LR to return to an upstream reaction zone of the flow point K, and the reaction product of the circulating liquid flow XK-LR flows through the point K and exists in XK. The mode of forming the back flow can be any suitable mode, such as arranging a built-in inner circulation tube, a built-in outer circulation tube, a built-in liquid collecting cup, a flow guide tube, a circulating pump, an external circulating tube and the like.

The invention discloses a liquid product circulating upflow type expanded bed hydrogenation reactor system, which is characterized in that a liquid product returns to an upstream reaction space for circular processing or liquid product circulation exists in an operation mode of a reaction zone or a main reaction zone of an expanded bed reactor; the liquid product circulation in the hydrogenation reactor refers to that at least a part of the liquid phase XK-L in the intermediate product XK or the final product XK at the flow point K is used as a circulating liquid flow XK-LR to return to a reaction area upstream of the flow XK, and the circulating liquid flow XK-LR passes through the point K and exists in XK. The way of forming the circulation of the liquid product can be any suitable way, but a gas-liquid separation zone must be arranged in the head space in the reactor to obtain the circulating liquid and other products, namely a built-in liquid collecting cup, a diversion pipe and a circulating booster, wherein the circulating booster is usually a circulating pump and can be arranged inside or outside the reactor.

The liquid collecting cup or the liquid collector arranged in the reactor refers to a container which is arranged in the reactor and is used for collecting liquid, the upper part or the upper part of the container is usually provided with an opening on the side surface, and a guide pipe is arranged on the bottom part or the lower part of the container for conveying or discharging the collected liquid; the top liquid collector of the expanded bed reactor is usually arranged in a liquid removal area of gas-liquid materials to obtain liquid and gas-liquid mixed phase material flow containing a small amount of bubbles or obtain liquid and gas, and at least part of liquid phase products are pressurized by a circulating pump and then return to a reaction space for circular processing. Typical examples are the heavy OIL ebullated bed hydrogenation reactor, the HTI coal hydrogenation direct liquefaction reactor used in the H-OIL process.

The thermal high separator refers to a gas-liquid separation device for separating intermediate products or final products of hydrogenation reaction.

The two-stage or multi-stage hydrogenation method of the invention refers to a hydrogenation method comprising two reaction stages or a plurality of reaction stages.

The hydrogenation reaction stage refers to a flow path section from the beginning of a hydrogenation reaction process of a hydrocarbon raw material to the gas-liquid separation of a hydrogenation product of the hydrocarbon raw material to obtain at least one liquid-phase product consisting of at least one part of generated oil, and comprises the hydrogenation reaction process of the hydrogenation reaction stage and the gas-liquid separation process of at least one part of the hydrogenation reaction product of the hydrogenation reaction stage. Therefore, the first-stage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material only comprises one hydrogenation reaction step and a gas-liquid separation process of a product of the hydrogenation reaction step, wherein 1 or 2 or more hydrogenation reactors which are operated in series can be used according to the requirement of the hydrogenation reaction step, so that the number and the form of the reactors are not the basis for determining the reaction level, and the reaction step consisting of one or a plurality of series reactors and the product separator are combined together to form the hydrogenation reaction level in the sense of completion.

The secondary hydrogenation method of the invention refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and is formed by two different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps, wherein at least a part of a flow formed by the oil generated by the primary hydrogenation enters the secondary hydrogenation reaction process.

The three-stage hydrogenation method refers to a flow mode that the processing process of an initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and is formed by three different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps, wherein at least one part of a material flow formed by the oil generated by the first-stage hydrogenation enters a second-stage hydrogenation reaction process, and at least one part of a material flow formed by the oil generated by the second-stage hydrogenation enters a third-stage hydrogenation reaction process. The flow structure of the hydrogenation method with more stages can be analogized according to the principle. The multistage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and consists of three or more different hydrogenation reaction processes and hydrogenation product gas-liquid separation processes.

The three-stage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and comprises three different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps.

The invention relates to a method similar to a two-stage hydrogenation method, which is a method similar to the two-stage hydrogenation method, and is regarded as the two-stage hydrogenation method when the ratio of the flow of a back-mixing liquid phase of a rear-stage upper feeding back-mixing flow expansion bed reactor to the flow of a liquid phase in an upper feeding tends to be infinite.

In the upflow hydrocracking reaction process R10 of the heavy oil UR10F, the reaction product BASE-R10P is at least a gas-liquid two-phase material flow, and in most cases, the material flow belongs to a gas-liquid-solid three-phase material flow. The hydrogenation reaction effluent R10P is used for discharging a hydrogenation reaction product BASE-R10P, appears in the form of 1-path or 2-path or multi-path materials, and is a gas phase or liquid phase or gas-liquid mixed phase or gas-liquid-solid three-phase material flow.

The upflow reactor of the invention can work in a mode selected from:

firstly, a suspension bed hydrogenation reactor;

a fluidized bed hydrogenation reactor, wherein the catalyst with reduced activity is usually discharged from the bottom of a bed layer in an intermittent mode, and fresh catalyst is supplemented from the upper part of the bed layer in an intermittent mode to maintain the catalyst inventory of the bed layer;

③ combined hydrogenation reactor of suspension bed and fluidized bed

Fourthly, micro-expansion bed.

The hydrocarbon material of the invention comprises hydrocarbon liquid material such as heavy oil and/or hydrocarbon powder material such as coal powder.

The naphtha component of the present invention refers to hydrocarbon components having a conventional boiling point in the range of from about five to about 180 c.

The diesel component refers to a hydrocarbon component with a conventional boiling point of 165-370 ℃.

The wax oil component refers to a hydrocarbon component with a conventional boiling point of 350-550 ℃.

The light wax oil component refers to a hydrocarbon component with a conventional boiling point of 350-425 ℃.

The medium-quality wax oil component refers to a hydrocarbon component with a conventional boiling point of 425-480 ℃.

The heavy wax oil component refers to a hydrocarbon component with a conventional boiling point of 480-550 ℃.

The vacuum residue component of the present invention refers to a hydrocarbon component having a conventional boiling point of greater than 530 ℃.

The heavy oil component of the present invention refers to a hydrocarbon component having a normal boiling point of greater than 350 ℃.

The naphtha in the invention refers to distillate oil composed of naphtha components.

The naphtha of the invention refers to distillate oil composed of diesel components.

The wax oil refers to distillate oil consisting of wax oil components.

The vacuum residue in the invention refers to distillate oil composed of vacuum residue components.

The heavy oil of the present invention means distillate oil composed of wax oil component and/or vacuum residue oil component.

In the hydrocracking process of the vacuum residue by using the boiling bed, an unconverted vacuum residue circulating operation mode can be adopted, the ratio of the weight flow rate of the unconverted vacuum residue in the circulating reaction to the weight flow rate of the fresh vacuum residue is the circulating ratio of the unconverted vacuum residue in the circulating reaction, the circulating ratio is usually 0.6-1.0 time or even more, and the higher the asphaltene content of the fresh vacuum residue is, the higher the circulating ratio of the unconverted vacuum residue in the circulating reaction is;

in the vacuum residue boiling bed hydrocracking process, a product heavy wax oil (with high aromatic hydrocarbon concentration) circulation operation mode (partial circulation) can be adopted, the ratio of the weight flow rate of the circulation reaction heavy wax oil to the weight flow rate of the fresh vacuum residue is the circulation ratio of the circulation reaction heavy wax oil, the circulation ratio is usually 0.15-0.3 times or even more, and the higher the asphaltene content of the fresh vacuum residue is, the higher the circulation ratio of the circulation reaction tail oil is, and the product heavy wax oil can be used as an asphalt solvent or a hydrogen supply agent for circulation.

In the vacuum residue suspension bed hydrocracking process, an unconverted vacuum residue circulating operation mode (partial circulation or large circulation) can be adopted, for example, a vacuum residue suspension bed hydrocracking process using a molybdenum-based catalyst is taken as an example, the ratio of the weight flow rate of the unconverted vacuum residue in the circulating reaction to the weight flow rate of the fresh vacuum residue is the circulating ratio of the unconverted vacuum residue in the circulating reaction, the circulating ratio is usually 0.6-1.0 times or even more, and the higher the asphaltene content of the fresh vacuum residue is, the higher the circulating ratio of the unconverted vacuum residue in the circulating reaction is;

in the vacuum residue suspension bed hydrocracking process, a product heavy wax oil circulation operation mode (partial circulation or full circulation) is generally adopted, for example, a vacuum residue suspension bed hydrocracking process using a molybdenum-based catalyst is taken, a ratio of a weight flow rate of a circulation reaction heavy wax oil to a weight flow rate of fresh vacuum residue is a circulation ratio of the circulation reaction heavy wax oil, the circulation ratio is generally 0.15-0.3 times or even more, and the higher the asphaltene content of the fresh vacuum residue is, the higher the circulation ratio of the circulation reaction tail oil is, and the circulation ratio is used as an asphalt solvent or a hydrogen supply agent for circulation.

The characteristic parts of the present invention are described below.

(4) in the combined vacuum distillation process U20:

①(97∶3)～(93∶7)；

②(93∶7)～(88∶12)；

③(88∶12)～(76∶24)。

③ the solid particulate hydrocarbon material R10FS comprises coal fines.

In the present invention, the operating conditions may be:

first, 90-97 wt%;

(ii) 80 to 90 wt%;

(iii) 70-80 wt%;

and 50-70 wt%.

①1.0～2.0；

②0.5～1.0；

③0.30～0.5；

①0.5～0.8；

②0.3～0.5；

③0.15～0.3。

firstly, a suspension bed reactor is a slurry bed reactor;

② a fluidized bed reactor;

③ a combined reactor of a suspension bed and a fluidized bed;

fourthly, micro-expansion bed.

Claims

1. The combined distillation method for generating oil by tail liquid circulating hydrocarbon material upflow type hydrogenation thermal cracking is characterized by comprising the following steps:

2. The method of claim 1, further comprising:

3. The method of claim 1, further comprising:

(4) in the combined vacuum distillation process U20:

4. The method of claim 1, further comprising:

(4) the operating conditions of the combined vacuum distillation process U20 were selected from 1 of the following:

5. The method of claim 1, further comprising:

6. The method of claim 1, further comprising:

(4) combining feed U20-F of vacuum distillation process U20 into first feed SF1 and second feed SF2, the ratio of the weight flow rate SF1-WR of first feed SF1 to the weight flow rate SF1-WR of second feed SF2 being K10;

k10 ═ SF1-WR)/(SF2-WR), K10 values were selected from one of the following:

①(97∶3)～(93∶7)；

②(93∶7)～(88∶12)；

③(88∶12)～(76∶24)。

7. the method of claim 1, further comprising:

8. The method of claim 1, further comprising:

(4) in the combined vacuum distillation process U20, a combined vacuum distillation of the first feed SF1 and the second feed SF2 is carried out using one combined vacuum column;

9. The method of claim 8, further comprising:

(4) in the combined vacuum distillation process U20, the lower part of the combined vacuum tower is provided with a flash evaporation zone, and the working mode is selected from 1 of the following modes:

the second feeding flash liquid SF2L-RL for the cyclic reaction is discharged into the first flash zone in a mode that the elevation of a discharge port of the second feeding flash liquid SF2L-RL for the cyclic reaction discharged from the second flash zone is higher than the operation elevation of a gas-liquid interface of a liquid phase gathering zone of the first flash zone; alternatively, the discharge of the second feed flash SF2L-RL from the recycle reaction into the first flash zone may be at a higher elevation than the discharge of the second feed flash SF2L-RL from the recycle reaction into the first flash zone at the inlet of the first feed SF 1.

10. The method of claim 1, further comprising:

(1) the hydrocarbon feed R10F is selected from 1 of the following:

③ the solid particulate hydrocarbon material R10FS comprises coal fines.

11. The method of claim 1, wherein:

12. The method of claim 1, further comprising:

(1) the hydrocarbon material upflow type hydrocracking reaction process R10 is a vacuum residuum hydrocracking reaction process, and the weight hydrocracking rate of the components with the conventional boiling point higher than 550 ℃ in the liquid raw material R10FL is selected from one of the following:

first, 90-97 wt%;

(ii) 80 to 90 wt%;

(iii) 70-80 wt%;

and 50-70 wt%.

13. The method of claim 1, further comprising:

(1) the hydrocarbon material upflow type hydrogenation thermal cracking reaction process R10 is a vacuum residue hydrogenation thermal cracking reaction process, and the liquid raw material R10FL comprises a fresh vacuum residue raw material R10FL-NEW and total circulating reaction tail liquid U20-RR from the step (4);

①1.0～2.0；

②0.5～1.0；

③0.30～0.5；

①0.5～0.8；

②0.3～0.5；

③0.15～0.3。

14. the method of claim 1, further comprising:

(1) the working mode of the hydrocarbon material hydrocracking reactor used in the hydrocarbon material upflow type hydrocracking reaction process R10 is selected from 1 or more of the following modes:

firstly, a suspension bed reactor is a slurry bed reactor;

② a fluidized bed reactor;

③ a combined reactor of a suspension bed and a fluidized bed;

fourthly, micro-expansion bed.

15. The method of claim 1, further comprising:

(4) in combined vacuum distillation process U20, ② in second flash section U22, second feed SF2 is subjected to a second flash process in a second flash vessel for flash separation into second feed flash gas SF2V and second feed flash liquid SF2L containing resid components;

a heavy wax oil removing condensation section U20KD is arranged in the second flash evaporator, the second feeding flash evaporation gas SF2V goes upward to pass through the heavy wax oil removing condensation section U20KD and is in countercurrent contact with the wax-containing oil cooling liquid U20QL entering from the upper part of the heavy wax oil removing condensation section U20KD for cooling, at least part of heavy wax oil components in the second feeding flash evaporation gas SF2V are condensed into heavy wax oil removing liquid U20-HVGO which is discharged out of the second flash evaporator and is not mixed with the second feeding flash evaporation liquid SF 2L;

the gas leaving the heavy wax oil removal condensation section U20KD is taken as heavy gas and enters the rectification section of the first vacuum distillation tower of the first vacuum distillation part U21 for recovery.