CN109705904B

CN109705904B - Processing method and processing system for hydrocarbon oil capable of producing ethylene and propylene in high yield

Info

Publication number: CN109705904B
Application number: CN201711014510.4A
Authority: CN
Inventors: 刘银亮; 谢朝钢; 张执刚; 魏晓丽; 张策; 陈昀; 陈学峰
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2017-10-25
Filing date: 2017-10-25
Publication date: 2020-12-04
Anticipated expiration: 2037-10-25
Also published as: CN109705904A

Abstract

The invention discloses a hydrocarbon oil processing method and a hydrocarbon oil processing system for producing ethylene and propylene in a large quantity. The method comprises a catalytic cracking step and an oil-gas separation post-treatment step, wherein the catalytic cracking step comprises the following steps: raw oil, recycled light diesel oil and a catalytic cracking catalyst are subjected to contact reaction in a first reaction zone to generate a first oil-gas mixture; the recycled light naphtha, the raffinate oil extracted from heavy naphtha aromatic hydrocarbon and a catalytic cracking catalyst are subjected to contact reaction in a second reaction zone to generate a second oil-gas mixture; then the second oil-gas mixture and recycled C4 olefin are subjected to contact reaction in a third reaction zone to produce a third oil-gas mixture; sending the first oil-gas mixture and the third oil-gas mixture into a settler for carrying out spent catalyst and oil-gas separation to obtain product oil gas; according to the method, the low-added-value products produced in the oil-gas separation process are recycled by a specific route, so that the yields of ethylene, propylene and aromatic hydrocarbon can be effectively improved, and further the economic benefit and the crude oil utilization rate of the original catalytic cracking process are improved.

Description

Processing method and processing system for hydrocarbon oil capable of producing ethylene and propylene in high yield

Technical Field

The invention relates to the field of producing ethylene and propylene by processing hydrocarbon oil, in particular to a hydrocarbon oil processing method and a hydrocarbon oil processing system for producing ethylene and propylene in a high yield.

Background

Ethylene is an important organic chemical basic raw material, and is mainly used for producing polyethylene, ethylene propylene rubber, polyvinyl chloride and the like. Ethylene is also one of the largest chemical products in the world, and accounts for more than 75% of the whole petrochemical product yield in the world; propylene is also an important organic chemical raw material, and is mainly used for preparing acrylonitrile, propylene oxide, acetone and the like. Ethylene and propylene are increasingly required as important chemical intermediates for the production of various important organic chemical raw materials, the formation of synthetic resins, synthetic rubbers, various fine chemicals, and the like.

The traditional route of preparing ethylene and propylene by steam cracking is adopted, the demand for chemical light hydrocarbons such as light hydrocarbons and naphtha is large, 70 ten thousand tons of chemical light oil are expected to be needed in 2020, while the crude oil in China is generally heavier, the chemical light oil is difficult to meet the requirements of ethylene and propylene raw materials, and under the condition of insufficient petroleum resources, the diversification of the steam cracking raw materials becomes the industrial development trend of ethylene and propylene. The steam cracking raw materials mainly comprise light hydrocarbon (such as ethane, propane and butane), naphtha, diesel oil, condensate oil and hydrogenated tail oil, wherein the mass fraction of the naphtha accounts for more than 50%, the ethylene yield of typical naphtha steam cracking is about 29-34%, the propylene yield is 13-16%, the propylene/ethylene ratio is about 0.5, and the lower propylene/ethylene output ratio is difficult to meet the current propylene requirement.

By adopting the petroleum hydrocarbon catalytic cracking technology and optimizing the technological parameters and the catalyst properties, the selectivity of the low-carbon olefin can be effectively improved, and the light aromatic hydrocarbon can be produced in a large amount. In order to improve the yield of light olefins, especially ethylene and propylene, and the yield of light aromatics, the prior art mainly comprises:

CN1234426A contacts heavy petroleum hydrocarbon and water vapor with catalyst in the lower part of the composite reactor comprising riser and dense-phase fluidized bed; the light petroleum hydrocarbon enters the upper part of the composite reactor, namely the bottom of the dense phase fluidized bed to contact with the catalyst.

CN1667089A makes the raw oil or part of hydrogenated cycle oil contact with regenerated catalytic cracking catalyst and water vapor in the catalytic cracking reactor for reaction.

CN101362669A is that hydrocarbons with different cracking performances contact with a catalytic cracking catalyst, the cracking reaction is carried out in a fluidized bed reactor to obtain target products of low-carbon olefin, aromatic hydrocarbon and a re-cracked raw material, the light aromatic hydrocarbon is hydrogenated and extracted, the cracked raw material is returned to a riser, wherein C2-C4 alkane enters steam cracking for further reaction, the yield of ethylene and propylene is more than 20 weight percent, and the co-production of aromatic hydrocarbons such as toluene, xylene and the like is carried out at the same time.

CN101362961A proposes that raw materials with the distillation range of 160-260 ℃ are contacted with a catalytic cracking catalyst, and cracking reaction is carried out in a fluidized bed reactor to obtain target products of low-carbon olefin and aromatic hydrocarbon, so that the yield and selectivity of ethylene and propylene are increased.

CN101469275B discloses a method for producing low carbon olefins from inferior crude oil, which is characterized in that crude oil with a total acid value of more than 0.5mgKOH/g and a characteristic factor K value of more than 12.1 after pretreatment is preheated is directly introduced into a catalytic cracking reactor to contact with a catalyst, and reacts under the catalytic cracking reaction condition, oil gas and spent catalyst after reaction are separated, the reaction oil gas is sent to a subsequent separation system to obtain products with different distillation ranges, and the spent catalyst is recycled after stripping and regeneration.

CN102337154A proposes that under the condition of catalytic cracking, hydrocarbon raw materials are contacted with a catalytic cracking catalyst in a composite reactor to obtain products such as low-carbon olefin and gasoline rich in light aromatics, and the products are further separated to obtain the light aromatics.

CN102344832A discloses a catalytic conversion method of petroleum hydrocarbon, which is characterized in that raw oil is contacted with a catalyst rich in mesoporous zeolite in a reactor to react, a spent catalyst and reaction oil gas are separated, the spent catalyst is returned to the reactor after being regenerated, the reaction oil gas is separated to obtain a product comprising ethylene, propylene, propane, butane and catalytic gasoline rich in aromatic hydrocarbon, wherein the catalytic gasoline is subjected to selective hydrogenation and then enters an aromatic hydrocarbon extraction device to be separated to obtain a target product light aromatic hydrocarbon, and raffinate oil, propane and butane are subjected to steam cracking to further produce ethylene and propylene.

From the above prior art, the development of the technology for producing low-carbon olefins by catalytic conversion of hydrocarbons mainly focuses on the catalytic cracking technology of heavy oil, and since the distillation range of heavy oil is wide, the hydrocarbon molecules are large, the product structure is complicated, and the yield of non-target products is high, in order to increase the yield of low-carbon olefins, a high reaction temperature is usually adopted, so that the yield of dry gas, especially the yield of methane, is greatly increased under the condition of increasing the yield of low-carbon olefins.

The catalytic cracking technology using light raw oil such as naphtha as raw material is in the research and development stage, in terms of world, ethylene production always uses light hydrocarbon and naphtha as main cracking raw materials, and the proportion of paraffin base and intermediate base raw materials in the cracking raw materials is very small, but the light hydrocarbon raw oil resource in China is limited, most of the raw oil belongs to heavy oil, the straight-run naphtha yield is very small, and the proportion of the wax oil raw material in ethylene production is forced to be higher.

In order to meet the increasing demand of low-carbon olefins, particularly propylene, and the demand of light aromatics, and to improve the utilization rate of resources such as heavy raw materials, e.g., normal-pressure wax oil, it is necessary to develop a catalytic conversion method for reducing the catalytic cracking reaction temperature and converting the wax oil fraction into ethylene and propylene to the maximum extent, thereby realizing the efficient utilization of petroleum resources.

Disclosure of Invention

The invention aims to overcome one of the problems in the prior art, and provides a hydrocarbon oil processing method and a hydrocarbon oil processing system for increasing the yield of ethylene and propylene, so as to increase the yield of ethylene, propylene and aromatic hydrocarbon (BTX) simultaneously.

In order to achieve the above object, a first aspect of the present invention provides a hydrocarbon oil processing method for increasing the production of ethylene and propylene, the method comprising a catalytic cracking step and an oil-gas separation post-treatment step, the catalytic cracking step comprising: raw oil, recycled light diesel oil and a catalytic cracking catalyst are subjected to contact reaction in a first reaction zone to generate a first oil-gas mixture; the recycled light naphtha, the raffinate oil extracted from heavy naphtha aromatic hydrocarbon and a catalytic cracking catalyst are subjected to contact reaction in a second reaction zone to generate a second oil-gas mixture; then the second oil-gas mixture and recycled C4 olefin are subjected to contact reaction in a third reaction zone to produce a third oil-gas mixture; sending the first oil-gas mixture and the third oil-gas mixture into a settler for carrying out spent catalyst and oil-gas separation to obtain product oil gas; in the oil-gas separation post-treatment step, the product oil gas is further divided into hydrogen, methane, products of ethylene and propylene, ethane, propane, C4 alkane, C4 olefin, light naphtha, heavy naphtha, light diesel oil, heavy diesel oil and the components, and the heavy naphtha is subjected to aromatic extraction to form light aromatic hydrocarbon and raffinate oil after heavy naphtha aromatic hydrocarbon extraction; wherein the C4 olefin and raffinate oil after extraction of light naphtha, light diesel oil and heavy naphtha aromatic hydrocarbon are refluxed to the corresponding reaction zone together.

Meanwhile, in a second aspect of the present invention, there is also provided a hydrocarbon oil processing system for the high yield of ethylene and propylene, the system comprising a catalytic cracking unit including: the first riser reactor is provided with a raw oil inlet, a light diesel oil refining port and a first reaction mixture outlet; the second riser reactor is provided with a raffinate oil recycling port after heavy naphtha aromatic hydrocarbon is extracted, a light naphtha recycling port and a third mixture outlet; a third riser reactor, wherein the third riser reactor is provided with a reaction mixture inlet, a C4 olefin recycling outlet and a third reaction mixture outlet which are connected with the second reaction mixture outlet; the sedimentation device is provided with a reaction mixture inlet, a spent agent outlet and a product oil gas outlet, wherein the reaction mixture inlet is respectively connected with the first reaction mixture outlet and the third reaction mixture outlet; the regeneration device is provided with a spent agent inlet connected with the spent agent outlet and a regenerated agent outlet, and the regenerated agent outlet is respectively connected with the catalyst inlets of the first riser reactor and the second riser reactor.

By applying the hydrocarbon oil processing method and the hydrocarbon oil processing system for producing ethylene and propylene in a large amount, the low-added-value products produced in the oil-gas separation process are recycled by a specific route, so that the yields of ethylene, propylene and aromatic hydrocarbon can be effectively improved, and the economic benefit and the crude oil utilization rate of the original catalytic cracking are further improved.

Drawings

FIG. 1 shows a schematic structural diagram of a riser reactor according to an embodiment of the invention;

FIG. 2 shows a schematic structural view of a catalytic cracking unit according to an embodiment of the present invention;

FIG. 3 is a schematic view showing the structure of a hydrocarbon oil processing system according to an embodiment of the present invention.

Description of the reference numerals

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

The invention provides a hydrocarbon oil processing method for producing ethylene and propylene in a high yield, which comprises a catalytic cracking step and an oil-gas separation post-treatment step, wherein the catalytic cracking step comprises the following steps: carrying out a first contact reaction on raw oil, recycled light diesel oil and a catalytic cracking catalyst in a first reaction zone to generate a first oil-gas mixture; carrying out a second contact reaction on the recycled light naphtha, the raffinate oil extracted from heavy naphtha aromatic hydrocarbon and a catalytic cracking catalyst in a second reaction zone to generate a second oil-gas mixture; then carrying out a third contact reaction on the second oil-gas mixture and recycled C4 olefin in a third reaction zone to produce a third oil-gas mixture; sending the first oil-gas mixture and the third oil-gas mixture into a settler for carrying out spent catalyst and oil-gas separation to obtain product oil gas; in the oil-gas separation post-treatment step, the product oil gas is further divided into hydrogen, methane, products of ethylene and propylene, ethane, propane, C4 alkane, C4 olefin, light naphtha, heavy naphtha, light diesel oil, heavy diesel oil and the components, and the heavy naphtha is subjected to aromatic extraction to form light aromatic hydrocarbon and raffinate oil after heavy naphtha aromatic hydrocarbon extraction; wherein the C4 olefin and raffinate oil after extraction of light naphtha, light diesel oil and heavy naphtha aromatic hydrocarbon are refluxed to the corresponding reaction zone together.

According to the process of the present invention, the aforementioned recycled raw materials can be partially or totally recycled according to the product requirements.

According to the method of the invention, an upper nozzle and a lower nozzle are included in the first reaction zone, in order to increase the conversion rate of the recycled light diesel oil, the raw oil is preferably fed from the upper nozzle, and the recycled light diesel oil is preferably fed from the lower nozzle;

preferably, in the interval of said first reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is between 0.2 and 5 seconds, preferably between 0.5 and 2 seconds, and may be, for example, any time value (in seconds) in the range of 0.5, 0.8, 1, 1.2, 1.5, 1.8, 2 and any two of these ratios. The ratio of solvent to oil is 10-100, preferably 30-75, and can be, for example, 30, 32, 35, 38, 42, 45, 48, 50, 52, 58, 62, 68, 70, 75, or any ratio in the range of any two of these ratios; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 0.8 to 9.8 seconds, preferably 2 to 6 seconds, and may be, for example, 2, 2.5, 2.8, 3.2, 3.5, 4.2, 4.8, 5.2, 5.8, 6, and any time value (in seconds) in the range of any two of these ratios may be 4 to 50, preferably 5 to 12, and may be, for example, 5, 6, 7, 8, 9, 10, 11, 12, and any ratio in the range of any two of these ratios.

According to the method of the invention, an upper nozzle and a lower nozzle are arranged in the second reaction zone, in order to increase the conversion rate of raffinate oil after heavy naphtha aromatic extraction, the light naphtha is preferably fed from the upper nozzle, and the raffinate oil after heavy naphtha aromatic extraction is preferably fed from the lower nozzle;

preferably, in the interval of the third reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is 0.2 to 5 seconds, preferably 0.5 to 3 seconds, and may be, for example, 0.5, 0.8, 1.2, 1.5, 2.2, 2.5, 2.8, 3 and any time value (seconds) in the range of any two of these ratios; the ratio of the solvent to the oil is 10-100, preferably 30-70, and can be, for example, 30, 35, 40, 45, 50, 55, 60, 65, 70 or any ratio in the range formed by any two of the ratios; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 0.8 to 9.8 seconds, preferably 1 to 8 seconds, and may be, for example, 0.8, 1.5, 2.0, 2.6, 3.2, 4.2, 5.6, 6.2, 7.3, 8 and any time value (seconds) in the range of any two of these ratios, and the agent-to-oil ratio is 5 to 60, preferably 15 to 30, and may be, for example, 15, 18, 20, 24, 26, 30 and any ratio in the range of any two of these ratios.

According to the process of the present invention, the residence time of the hydrocarbon oil in said third reaction zone is between 2 and 30 seconds, preferably between 6 and 15 seconds. No catalyst is separately added to the second reaction zone, wherein the catalyst is carried into the second reaction zone by the hydrocarbon oil in the first reaction zone. Wherein the C4 olefins are reclaimed at 10-100 wt% of the total C4 olefin product, wherein the hydrocarbon oil ratio is in the range of 3-50, preferably 4-12, and can be, for example, 4, 5, 6, 7, 8, 9, 10, 11, 12, or any ratio in the range of any two of these ratios.

According to the process of the present invention, the hydrocarbon oil is introduced from the lower part of the reaction zone and flows upward, wherein the material fed from the "lower nozzle" is introduced into the reaction zone first, and the material fed from the "upper nozzle" is introduced into the reaction zone after the material is introduced, and reacts with the fed material in a mixed manner.

According to the method of the present invention, the aforementioned "agent-oil ratio" refers to the weight ratio of the catalyst to the hydrocarbon oil, wherein the hydrocarbon oil refers to the total amount of the raw hydrocarbon oil entering the corresponding "interval", for example, the hydrocarbon oil is the amount of the recycled light diesel oil in the "interval between the upper nozzle and the lower nozzle" of the first reaction zone, and the hydrocarbon oil is the total amount of the raw oil and the recycled light diesel oil in the "interval between the upper nozzle and the outlet".

The process according to the invention, wherein the conditions for the first, second and third reaction zones can be selected within a relatively wide range; preferably, the contact reaction conditions of the first reaction zone include: the reaction temperature is 500-750 ℃, preferably 520-590 ℃, and can be, for example, 530 ℃, 540 ℃, 550 ℃, 560 ℃, 570 ℃, 580 ℃, 590 ℃ or any value in the range formed by any two of the values; the residence time is 1 to 50 seconds, preferably 1 to 10 seconds; the pressure is 0.1-1.0MPa, preferably 0.2-0.3 MPa; the agent-oil ratio is 4-40, preferably 5-12, and can be, for example, any ratio in the range of 5, 6, 7, 8, 9, 10, 11, 12 and any two of the ratios; the weight ratio of the pre-lifting medium to the raw oil is 0.03-1.0, preferably 0.05-0.20;

preferably, the contact reaction operating conditions of the second reaction zone include: the reaction temperature is 500-; the residence time is 1 to 50 seconds, preferably 1 to 10 seconds; the pressure is 0.1-1.0MPa, preferably 0.2-0.3MPa, the agent-oil ratio is 5-60, preferably 10-40, and can be, for example, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, and any ratio in the range formed by any two of the ratios; the weight ratio of the pre-lifting medium to the raw oil is 0.03-1.0, preferably 0.03-0.05.

Preferably, the contact reaction conditions of the third reaction zone include: the reaction temperature is 490-740 ℃, preferably 600-660 ℃, and can be, for example, 600 ℃, 610 ℃, 620 ℃, 630 ℃, 640 ℃, 650 ℃, 660 ℃ or any value in the range formed by any two of these values; the residence time is 1 to 50 seconds, preferably 2 to 20 seconds; the pressure is 0.1-1.0MPa, preferably 0.2-0.3 MPa; the ratio of solvent to oil is 1 to 100, preferably 15 to 25, and may be, for example, any ratio in the range of 15, 16, 17, 18, 19, 20, 21, 22, 23, 24 and 25 and any two of these ratios.

According to the method of the present invention, the aforementioned "agent-oil ratio" refers to the weight ratio of the catalyst to the hydrocarbon oil, wherein the hydrocarbon oil refers to the sum of the raw materials to be reacted entering the reaction zone, for example, the hydrocarbon oil is the sum of the raw material oil and the raffinate oil after the optional subsequent heavy naphtha aromatic extraction in the first reaction zone, and the hydrocarbon oil is the sum of the light diesel oil in the second reaction zone.

According to the method, pre-lifting media are introduced into the first reaction zone and the second reaction zone to drive the catalytic cracking catalyst to rise, so that plug flow of the catalytic cracking catalyst with uniform density is formed at the bottom of the reactor. Wherein the pre-lifting medium can be one or more of steam, refinery dry gas, light alkane and light olefin.

The process according to the invention, wherein the three contact reactions can be carried out in any catalytic cracking reactor, for example both in a riser reactor and in a fluidized bed reactor; the method can be carried out in an up-flow reactor or a down-flow reactor; preferably, the three contact reactions are all carried out in a riser reactor, and when the three contact reactions are carried out in the riser reactor, the catalytic cracking device self dry gas is preferably used for replacing pre-lifting steam, and is introduced into the reactor after partial/total replacement of atomized steam, and the dry gas yield is reduced by increasing the hydrogen partial pressure of the dry gas, so that the selectivity of propylene can be further improved.

The method according to the present invention focuses on the improvement of the process flow, wherein there may be no particular requirement for the catalytic cracking catalyst used in the catalytic cracking process as long as it is suitable for use as a catalytic cracking catalyst, preferably a catalytic cracking catalyst containing a medium pore zeolite; more preferably, the catalytic cracking catalyst comprises, based on the total weight of the catalyst, from 2 to 60 wt% of zeolite, from 10 to 99 wt% of an inorganic oxide, from 0 to 70 wt% of clay; wherein the zeolite comprises a medium pore zeolite and optionally a large pore zeolite, preferably the zeolite comprises from 50 to 95 wt% of the medium pore zeolite and from 5 to 50 wt% of the large pore zeolite, more preferably the zeolite comprises from 70 to 95 wt% of the medium pore zeolite and from 5 to 30 wt% of the large pore zeolite, based on the total weight of the zeolite.

In the catalytic cracking catalyst used in the present invention, the medium pore zeolite and the large pore zeolite are defined as usual in the art, i.e., the medium pore zeolite has an average pore diameter of 0.5 to 0.6nm and the large pore zeolite has an average pore diameter of 0.7 to 1.0 nm. The medium pore zeolites which may be selected for use in the present invention are preferably zeolites having the MFI structure, such as ZSM series zeolites and/or ZRP zeolites; preferably, the mesoporous zeolite used in the present invention may be a mesoporous zeolite modified with a nonmetallic element such as phosphorus and/or a transition metal element such as iron, cobalt, nickel; for a detailed description of ZRP zeolites, reference may be made, inter alia, to the description in US5,232,675, for example ZRP-1; wherein the ZSM series zeolite is one or a mixture of more selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolite with similar structure, wherein the detailed description of the ZSM-5 can be seen in U.S. Pat. No. 3,702,886. The large pore zeolite may be selected from a mixture of one or more of the group of zeolites consisting of rare earth Y, rare earth hydrogen Y, ultrastable Y obtained by different processes, high silicon Y.

In the catalytic cracking catalyst used in the present invention, the inorganic oxide mainly functions as a binder, and is preferably selected from silicon dioxide and/or aluminum oxide. Wherein the silicon dioxide precursor is a substance which can be converted into silicon dioxide under the roasting condition, and preferably can be silica sol; wherein the alumina precursor is a substance which can be converted into alumina under the roasting condition, and preferably hydrated alumina and/or alumina sol; the hydrated alumina is one or more selected from boehmite, pseudo-boehmite, alumina trihydrate and amorphous aluminum hydroxide.

In the catalytic cracking catalyst used in the invention, the clay mainly plays a role of a matrix, preferably the clay is one or more selected from diatomite, expanded perlite, kaolin, halloysite, silicalite, hydrolyzed silica, macroporous silica and silica gel, and preferably selected from kaolin and/or halloysite.

According to the method of the invention, the catalytic cracking catalyst forms a spent catalyst after undergoing a catalytic cracking reaction, the spent catalyst forms a regenerant through regeneration, the formed regenerant is conveyed to the first reaction zone and the second reaction zone for reuse, wherein the conventional regeneration conditions of the corresponding catalyst are adopted for the process conditions of the regeneration treatment.

According to the method of the invention, in order to further increase the yield of ethylene and propylene products, preferably, the oil-gas separation post-treatment step further comprises the step of carrying out contact reaction on the ethane, propane and C4 alkane with water vapor to generate ethylene and propylene products and by-producing ethane, propane, C4 alkane, C4 alkene and steam cracking light naphtha, wherein the contact temperature is preferably 700-1000 ℃; preferably, the by-product ethane, propane and C4 alkane are refluxed and continuously contacted with water vapor for reaction; the C4 olefins and the steam cracked light naphtha that are by-produced are preferably returned to the third reaction zone.

In the process according to the present invention, in the step of subjecting the heavy naphtha to aromatic extraction to form light aromatics and raffinate oil after heavy naphtha aromatic extraction, there may be no limitation on the extraction solvent, for example, the extraction solvent is selected from one or more of sulfolane, N-methylpyrrolidone, diethylene glycol ether, triethylene glycol ether, tetraethylene glycol, dimethyl sulfoxide and N-formyl morpholine ether; there may be no particular requirement for the operating conditions of the aromatic extraction, for example the operating conditions of the aromatic extraction include: the temperature is 40-120 ℃, the volume ratio of the extraction solvent to the heavy naphtha is 0.5-5: 1.

according to the method of the present invention, preferably, the raw oil is heavy oil; more preferably, the raw oil is a hydrotreated heavy wax oil, preferably a hydrotreated heavy wax oil with a flow path of more than 350 ℃.

According to the process of the present invention, preferably, the heavy oil as feedstock is selected from or comprises petroleum hydrocarbons andand/or other mineral oils. The petroleum hydrocarbon is selected from one or more of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, atmospheric residual oil, hydrotreating gas oil, catalytic diesel oil and hydrogenation catalytic diesel oil, and the other mineral oil is selected from one or more of coal liquefaction oil, oil sand oil and shale oil. Preferably, the heavy oil is a paraffinic or a mid-based feedstock, preferably the heavy oil has a density of 0.87-0.93g/cm³. Preferably 0.88 to 0.91g/cm³. The carbon residue value is from 0 to 6% by weight, preferably from 0 to 2% by weight.

According to the process of the present invention, the main operating conditions of the hydrotreatment preferably include: the reaction temperature is 350-400 ℃, preferably 360-390 ℃, and can be, for example, 360 ℃, 370 ℃, 380 ℃, 390 ℃ or any value in the range formed by any two of the values; the reaction pressure is 7-15MPa, preferably 9-13MPa, and the volume space velocity is 0.5-1.5h^-1Preferably 0.6 to 1.2h^-1The hydrogen-oil ratio is 400-1500, preferably 600-1200.

According to the method of the present invention, preferably, the hydrogenation catalyst used in the hydrogenation treatment comprises: the hydrogenation catalyst comprises a carrier and a VIB group metal and/or VIII group non-noble metal oxide which are loaded on the carrier, wherein the content of the VIB group metal in an oxidation state is 3-35 wt%, and the content of the VIII group non-noble metal in the oxidation state is 0.3-8 wt%; preferably, the carrier is an alumina and/or silica-alumina carrier; preferably the group VIB metal is Mo and/or W; preferably the group VIII metal is Co and/or Ni. Wherein the silicon-aluminum carrier refers to a composite carrier containing aluminum oxide and silicon oxide.

According to the method of the invention, in order to further reduce low added value products and increase the yield of ethylene, propylene and aromatic hydrocarbons without increasing energy consumption, the heavy diesel oil and the components are preferably refluxed to the step of hydrotreating, and preferably in actual operation, the slurry of the lower layer of the heavy diesel oil and the components is discharged, and the upper layer of oil is refluxed to the step of hydrotreating.

Meanwhile, corresponding to the above-mentioned hydrocarbon oil processing method for producing ethylene and propylene in high yield provided by the present application, the present invention also provides a hydrocarbon oil processing system for producing ethylene and propylene in high yield, the system includes a catalytic cracking unit 30, as shown in fig. 1 and 2, the catalytic cracking unit 30 includes: the system comprises a first riser reactor A, a second riser reactor A and a third riser reactor B, wherein a raw oil inlet, a light diesel oil refining port and a first reaction mixture outlet are formed in the first riser reactor A; the second riser reactor B is provided with a raffinate oil recycling port after heavy naphtha aromatic hydrocarbon is extracted, a light naphtha recycling port and a third mixture outlet; a third riser reactor C, wherein a reaction mixture inlet, a C4 olefin recycling outlet and a third reaction mixture outlet which are connected with the second reaction mixture outlet are arranged on the third riser reactor C; the settling device 32 is provided with a reaction mixture inlet, a spent agent outlet and a product oil gas outlet, wherein the reaction mixture inlet is respectively connected with the first reaction mixture outlet and the third reaction mixture outlet; the regeneration device 33 is provided with a spent agent inlet connected with the spent agent outlet and a regenerated agent outlet, and the regenerated agent outlet is respectively connected with the catalyst inlets of the first riser reactor A and the second riser reactor B.

The system according to the present invention, wherein the settling device 32 and the regenerating device 33 are conventional devices, and the settling device 32 comprises a cyclone 321 and a stripper 322; the regeneration device 33 comprises a cyclone 331 and a main air distribution plate 332; a spent agent circulating slide valve 323 is formed on a spent agent conveying pipeline between the settling device 32 and the regeneration device 33; a regenerant circulating slide valve 333 is respectively arranged on a regenerant conveying pipeline between the regenerating device 33 and the first riser reactor a and the second riser reactor B.

According to the system of the present invention, preferably, an upper nozzle and a lower nozzle are arranged inside the first riser reactor a, the upper nozzle is connected with the raw oil inlet, and the lower nozzle is connected with the light diesel oil recycle port.

According to the system of the present invention, preferably, an upper nozzle and a lower nozzle are arranged inside the third riser reactor B, the upper nozzle is connected to the light naphtha recycle port, and the lower nozzle is connected to the raffinate oil recycle port after heavy naphtha aromatic extraction.

According to the system of the present invention, preferably, the system further comprises a raw oil supply unit, wherein the raw oil supply unit is connected with the raw oil inlet of the first riser reactor a in a matching manner, and is used for supplying raw oil to the first riser reactor a;

the system according to the present invention, as shown in fig. 3, preferably further comprises a catalytic cracking fractionation unit 40, a gas separation unit, a light and heavy gasoline separation unit 50, and a light and heavy diesel separation unit 60; the catalytic cracking fractionation unit 40 is connected with the settling device 32 in a matching manner and is used for fractionating the product oil gas separated by the settling device into dry gas 42a, liquefied gas 42b, naphtha 41b, naphtha 42c, diesel oil and the above components 41 c; the gas separation unit 50 is cooperatively connected with the catalytic cracking fractionation unit 40, and is used for separating the dry gas 42a and the liquefied gas 42b to obtain hydrogen, methane, product ethylene and propylene, ethane, propane, C4 alkane and C4 alkene, wherein the C4 alkene flows back to the third riser reactor C; the light and heavy naphtha separation unit 50 is cooperatively connected with the catalytic cracking fractionation unit 40, and is configured to separate the naphthas 41B and 42c into a light naphtha 51a and a heavy naphtha 51B, wherein the light naphtha 51a is refluxed to the second riser reactor B; the light diesel oil and heavy diesel oil separation unit 60 is connected with the catalytic cracking fractionation unit 40 in a matching manner, and is used for separating the diesel oil and the components 41c into light diesel oil 61a and heavy diesel oil and the components 61b, wherein the light diesel oil 61a is refluxed to the first riser reactor a.

The system of the invention, wherein the catalytic cracking fractionating unit 40 is mainly used for fractionating catalytic cracking product oil gas, and comprises a catalytic cracking main fractionating tower 41 and a light component separation tank 42; wherein, the product oil gas obtained by the separation of the settling device is fractionated into tower top oil gas 41a, naphtha 41b, diesel oil and the above components 41c by the catalytic cracking main fractionating tower 41. The light component separation tank 42 includes an absorption tower, a desorption tower, a reabsorption tower, a stabilizer tower, and the like, and is capable of separating the overhead oil gas 41a discharged from the catalytic cracking main fractionation tower 41 into dry gas 42a, liquefied gas 42b, and naphtha 42c, wherein the dry gas 42a, the liquefied gas 42b are sent to a subsequent gas separation unit, and the naphtha 42c is sent to a light-heavy gasoline separation unit.

The system of the invention is mainly used for separating and treating the low-carbon olefin in the dry gas 42a and the liquefied gas 42b obtained by fractionation in the catalytic cracking fractionation unit 40. The gas separation unit comprises a rich gas compressor, a depropanizing tower, a debutanizing tower, a demethanizing tower, a deethanizing tower, an ethylene rectifying tower, a propylene rectifying tower, a butadiene rectifying tower and the like. In practical operation, dry gas and liquefied gas obtained by fractionation in the catalytic cracking fractionation unit 40 are sent to the depropanizer after being pressurized by the rich gas compressor. After separation, C4 and C5 fractions are discharged from the bottom of the depropanizing tower, and enter a debutanizing tower after cooling, the material flow at the top of the debutanizing tower enters a butene rectifying tower, the material flow at the bottom of the debutanizing tower discharges C5 fractions, and the C5 fraction can be returned to the second riser reactor B as light naphtha after being in cocurrent with the light naphtha; c4 olefins discharged from the top of the butene rectification column are returned to the third riser reactor C in the catalytic cracking unit 30, and C4 alkanes at the bottom of the butene rectification column are optionally returned to the catalytic cracking unit 30 or conveyed to a subsequent distillation cracking unit. Depropanizer overhead C3^—Cooling the distillate, sending the cooled distillate into a demethanizer, discharging fuel gas mainly containing methane and hydrogen from the top of the demethanizer, and sending the bottom stream of the demethanizer into a deethanizer; feeding the top material flow of the deethanizer into an ethylene rectifying tower, feeding the bottom material flow of the deethanizer into a propylene rectifying tower, discharging polymerization-grade ethylene from the top of the ethylene rectifying tower, and subsequently feeding the ethane discharged from the bottom of the ethylene rectifying tower into a steam cracking unit. The polymerization-grade propylene is discharged from the top of the propylene rectifying tower, and the propane discharged from the bottom of the propylene rectifying tower is sent to a cracking unit.

According to the system of the invention, in order to further increase the yield of the ethylene and the propylene, preferably, the system further comprises a distillation cracking unit which is matched and connected with the gas separation unit and is used for carrying out contact reaction on the ethane, the propane and the C4 alkane with water vapor so as to generate the ethylene and the propylene, and by-products of the ethane, the propane, the C4 alkene and the C4 alkane, C5 components and steam cracking light naphtha, wherein the by-products of the ethane, the propane and the C4 alkane are refluxed to the distillation cracking unit; refluxing the co-produced C4 olefin to the third riser reactor C; and refluxing the byproduct steam cracked light naphtha to the second riser reactor B. The reaction conditions of the distillative cracking unit are as described in the previous method section.

According to the system of the present invention, preferably, the system further includes an aromatic hydrocarbon extraction unit, the aromatic hydrocarbon extraction unit is connected to the light and heavy naphtha separation unit 50 in a matching manner, and is configured to perform aromatic hydrocarbon extraction on heavy naphtha and obtain raffinate oil after heavy naphtha aromatic hydrocarbon extraction, and the raffinate oil after heavy naphtha aromatic hydrocarbon extraction may flow back to the second riser reactor B, or may be conveyed to a steam cracking unit for cracking. The aromatic hydrocarbon extraction unit can adopt a conventional aromatic hydrocarbon extraction unit, and for example, the aromatic hydrocarbon extraction unit can comprise an extractive distillation tower, an extractive distillation solvent recovery tower, a liquid-liquid extraction tower, a stripping tower, a liquid-liquid extraction recovery tower and the like. The reaction conditions of the aromatics extraction unit and the extraction solvent used are as described in the previous method section.

According to the system of the present invention, preferably, the system further comprises a hydrofining unit 10 and a hydro-fractionation unit 20, wherein the hydrofining unit 10 is cooperatively connected with the raw oil supply unit and is used for hydrotreating the raw oil to obtain a hydrogenated product; the hydro-fractionation unit 20 is connected with the hydrofining unit 10 in a matching manner and is used for fractionating the hydrogenation product to obtain hydrotreated heavy wax oil; the hydro-fractionation unit 20 is connected to the raw oil inlet of the first riser reactor a in the catalytic cracking unit 30 in a matching manner, and is configured to supply the hydrotreated heavy wax oil to the first riser reactor a.

The system according to the present invention, wherein the hydroprocessing unit 10 can be a conventional hydroprocessing unit, for example, mainly comprises therein a fresh hydrogen compressor 11, a feedstock preheater 12, a hydrofinishing reactor 13, a heat exchanger 14, a water cooler 15, a high pressure separator 16 and a low pressure separator 17. Wherein the new hydrogen compressor 11 is connected with the hydrofining reactor 13 through the raw material preheating furnace 12, the outlet end of the hydrofining reactor 13 is connected with the subsequent hydrofining unit through a high-pressure separator 16 and a low-pressure separator 17 which are arranged in sequence, wherein the high-pressure separator 16 separates hydrogen mixed in the hydrofining product, the part of circulating hydrogen returns to the raw material preheating furnace 12, and the low-pressure separator 17 removes the overhead gas in the hydrofining product. In addition, in order to form a heat balance of the entire system and to be suitable for the temperature of the high-pressure separator 16, it is preferable that a heat exchanger 14 and a water cooler 15 are further formed between the hydrofining reactor 13 and the high-pressure separator 16. The reaction conditions of the hydrofinishing reactor 13 are as described in the previous process section.

The system according to the present invention, wherein the hydro-fractionation unit 20 is mainly used for fractionating the hydrofined products, which includes the hydrofined fractionating tower 21, and the hydrofined products are fractionated into the overhead non-condensable gas 21a, the naphtha fraction 21b, the aviation kerosene fraction 21c, the high quality vehicle diesel oil fraction 21d and the bottom heavy wax oil (hydrotreated heavy wax oil, preferably heavy components with a distillation range of more than 350 ℃) through the hydrofined fractionating tower 21, wherein the bottom heavy wax oil is sent to the subsequent catalytic cracking unit.

According to the system of the present invention, in order to further improve the utilization rate of the catalytic cracking byproducts and reduce the low value added products, the heavy diesel oil and the above components separated by the light and heavy diesel oil separation unit 70 are preferably returned to the hydrorefining unit.

The method and system for processing a hydrocarbon oil with high yields of ethylene and propylene according to the present invention will be further described with reference to the following specific examples and comparative examples.

The heavy raw oil 1 used in the following examples and comparative examples was hydrogenated wax oil derived from the marine products of the company Zhenhai, Inc. of petrochemical industries, China, and the heavy raw oil 2 was vacuum wax oil derived from the company Daqing petrochemical products, Inc. of petroleum and gas, China, and its main properties are shown in Table 1.

Table 1.

Item	Heavy feedstock 1	Heavy feedstock 2
			Density (20 ℃, g/cm)³	0.903	0.8582
Kinematic viscosity (100 ℃ C.), mm²/s	8.325	4.174
			Freezing point, deg.C	-15	-20
Kinematic viscosity (80 ℃ C.), mm²/s	12.582	6.143
			Aniline point, deg.C	-	103.4
Refractive index, (nD70)	-	1.4585
			Carbon residue in wt%	0.52	0.21
Basic nitrogen, ppm	236	155
			C，	84.53	85.98
N，ppm	582	417
			H，	12.56	13.66
S,ppm	1352	885
			Initial boiling point	268	262
5％	321	310
			10％	345	333
30％	395	382
			50％	523	414
70％	451	444
			90％	502	490
95％	526	513
			Paraffin content, wt.%	28	14

The preparation methods of the catalytic cracking catalysts used in the following examples and comparative examples are briefly as follows:

1) 20gNH₄Cl is dissolved in 1000g of water, 100g of crystallized product ZRP-1 zeolite is added to the solution, the rare earth content RE₂O₃2.0 wt%), exchanged at 90 ℃ for 0.5h, filtered to obtain a filter cake; 4.0gH was added₃PO₄With 4.5gFe (NO)₃)₃Dissolving in 90g of water, mixing with a filter cake, soaking and drying; followed by calcination treatment at 550 ℃ for 2 hours to obtain a phosphorus and iron containing MFI structured mesoporous zeolite comprising, by elemental analytical chemistry: 0.1 wt% of Na₂O, 5.1 wt% of Al₂O₃2.4 wt% of P₂O₅1.5 wt% of Fe₂O₃3.8 wt% of RE₂O₃And 88.1 wt% SiO₂。

2) Pulping 75.4kg of halloysite by 250kg of decationized water, adding 54.8kg of pseudo-boehmite, adjusting the pH to 2-4 by hydrochloric acid, uniformly stirring, standing and aging at 60-70 ℃ for 1 hour, keeping the pH at 2-4, reducing the temperature to below 60 ℃, adding 41.5kg of alumina sol, and stirring for 40 minutes to obtain mixed slurry.

3) Adding the MFI structure mesoporous zeolite containing phosphorus and iron prepared in the step 1) and DASY zeolite into the mixed slurry obtained in the step 2), uniformly stirring, spray-drying and forming, washing with ammonium dihydrogen phosphate solution, and washing to remove free Na⁺And drying to obtain a catalytic cracking catalyst sample, wherein the catalyst comprises 18 wt% of phosphorus and iron-containing MFI structure mesoporous zeolite, 2 wt% of DASY zeolite, 36 wt% of alumina and 44 wt% of kaolin.

The main properties of the catalytic cracking catalyst prepared as described above are shown in table 2.

Table 2.

The preparation of the hydrotreating catalyst used in the following examples and comparative examples is briefly as follows:

ammonium metatungstate and nickel nitrate were weighed and made up into 200mL of solution with water. The solution was added to 50g of alumina carrier, immersed at room temperature for 3 hours, and the immersion liquid was treated with ultrasonic waves for 30 minutes during the immersion, cooled, filtered, and dried in a microwave oven for about 15 minutes. The catalyst comprises the following components: 30.0% by weight of WO₃3.1 wt% NiO and 66.9 wt% alumina.

The extraction solvent used in the following examples and comparative examples was sulfolane, chemically pure.

Example 1

The adopted hydrocarbon oil processing system comprises: as shown in fig. 3;

the adopted hydrocarbon oil processing method comprises the following steps:

the heavy oil feedstock 1 (properties shown as sample A in Table 1) was fed to a feedstock preheater 12, mixed with hydrogen from a fresh hydrogen compressor 11, heated and fed to a hydrofinishing reactor 13 (reaction conditions included: temperature 370 ℃, hydrogen partial pressure 10MPa, hydrogen to oil (volume) ratio 800, and volumetric space velocity 1.0h^-1Wherein the hydrogenation catalyst is the hydrogenation catalyst provided above). The hydrorefiningThe reactor 13 is divided into three reaction zones, the intermediate zone being fed with cold hydrogen to adjust the reaction temperature. The reacted oil gas enters a heat exchanger 14, the materials after heat exchange enter a water cooler 15 for further cooling, then enter a high-pressure separator 16 for separating out recycle hydrogen (the recycle hydrogen is compressed by a recycle hydrogen compressor and then enters a raw material preheating furnace 2), heavy components separated by the high-pressure separator 16 enter a low-pressure separator 17 and then enter a hydrofining device fractionating tower 21, non-condensable gas 21a at the top of the tower is separated out at the top of the tower and enters a gas pipe network, naphtha fraction 21b is extracted from the middle section of the fractionating tower, and the aviation kerosene fraction 21c and heavy components (hydrotreating heavy wax oil) with the flow of high-quality vehicle diesel oil being more than 350 ℃ and the flow of the bottom of the tower are extracted from the middle.

Preheating the hydrotreated heavy wax oil to 260 ℃ as raw oil, feeding the raw oil into a first riser reactor A (a first reaction zone) in a catalytic cracking unit through a riser pipeline, and carrying out contact reaction (the reaction conditions refer to table 3) on the hydrotreated heavy wax oil, the recycled light diesel oil (LCO)31a (LCO enters a lower nozzle, and the hydrotreated heavy wax oil enters an upper nozzle), atomized steam and a catalytic cracking catalyst (the catalytic cracking catalyst provided in the previous step) to produce a first reaction mixture; at the same time, the recycled heavy naphtha aromatic extracted Raffinate (RAA)31B and the recycled light naphtha 31C (lcn) are respectively fed into a second riser reactor B (second reaction zone, in which the light naphtha is fed into an upper nozzle and the heavy naphtha aromatic extracted raffinate is fed into a lower nozzle) in the catalytic cracking unit through a return line, and are subjected to a contact reaction with atomized steam and a catalytic cracking catalyst (provided as above) (see table 3 for reaction conditions), so as to generate a second reaction mixture, the second reaction mixture is fed into a third riser reactor C (third reaction zone) in the catalytic cracking unit, and is subjected to a third contact reaction with recycled C4 olefin 31d (see table 3 for reaction conditions), so as to generate a third reaction mixture, and the first reaction mixture and the third reaction mixture are fed into a cyclone 321 of a settling device 32 (with a pressure of 0.13MPa) together to separate a to-be-generated agent and produce a product The product oil gas and spent agent enter a stripper 322 for stripping, and the stripped spent agent enters a regenerator 33 through a spent agent circulating slide valve 323 to contact with air entering a main air distribution plate 332 at the bottom of the regenerator (pressure)0.15MPa, 700 ℃ and a bed density of 3kg/m³) Burning; the produced flue gas and the regenerant are separated by the cyclone 331, the flue gas enters the flue, the regenerant enters the bottom of the regenerator, and the flue gas and the regenerant enter the first riser reactor A and the second riser reactor B in the catalytic cracking unit through the regenerant circulating slide valve 333.

The product oil gas separated by the settling device 32 enters the catalytic cracking main fractionating tower 41 through an oil gas conveying pipeline and is divided into tower top oil gas 41a, naphtha 41b, diesel oil and the above 41 c. Wherein the oil gas 41a at the top of the tower enters a light component separation tank 42 through an oil transfer line, and the oil gas 41a at the top of the tower is separated into dry gas (mainly H) through an absorption tower, a desorption tower, a reabsorption tower, a stabilization tower and the like₂、CH₄、C₂Component) 42a, liquefied gas (C3, C4)42b, and naphtha 42C;

dry gas 42a and liquefied gas 42b generated by separating the light component separation tank 42 enter a gas separation unit (comprising a rich gas compressor, a depropanizing tower, a debutanizing tower, a demethanizing tower, a deethanizing tower, an ethylene rectifying tower, a propylene rectifying tower, a butadiene-lean rectifying tower and the like) through an oil gas conveying pipeline, and are separated into hydrogen, methane, products of ethylene and propylene, ethane, propane, C4 alkane, C4 alkene and C5 fractions; wherein the C4 olefins are refluxed to the third riser reactor C for recycle;

conveying ethane, propane and butane produced by the gas separation unit to a steam cracking unit to contact and react with steam at 830 ℃ to generate products of ethylene and propylene, ethane, propane, C4 olefin and C4 alkane, C5 component and steam cracking light naphtha, wherein the ethane, propane and C4 alkane return to the steam cracking unit to produce ethylene and propylene, and the C4 olefin returns to the third riser reactor C to be recycled; and refluxing the steam cracked light naphtha to the second riser reactor B for recycling;

naphtha 41b produced by a catalytic cracking main fractionating tower 41 and naphtha 42c separated by a light component separation tank 42 are conveyed to a light and heavy naphtha separation unit 50 through an oil gas pipeline and separated into light naphtha 51a and heavy naphtha 51 b; wherein the light naphtha 51a is refluxed to the second riser reactor B for recycle;

conveying the heavy naphtha 51b separated in the light and heavy naphtha separation unit 50 to an aromatic hydrocarbon extraction unit, extracting by sulfolane, wherein the extraction temperature is 100 ℃, and the volume ratio of the sulfolane solvent to the gasoline is 3.0: 1, separating out light aromatic hydrocarbon (BTX) and raffinate oil extracted from catalytic heavy naphtha aromatic hydrocarbon, wherein the raffinate oil extracted from the catalytic heavy naphtha aromatic hydrocarbon reflows to a second riser reaction B for remilling;

conveying diesel oil generated by a catalytic cracking main fractionating tower 41 and the components 41C to a light diesel oil separation unit 60 through an oil-gas pipeline, separating the diesel oil into light diesel oil 61a and heavy diesel oil and components 61b, wherein the light diesel oil 61a flows back to a first riser reactor A for recycling; the heavy diesel oil and the above components 61b are returned to the hydrorefining reactor 13 for recycling.

The product distribution after stable operation of the apparatus is shown in table 4.

Examples 2 to 3

The adopted hydrocarbon oil processing system comprises: referring to FIG. 3, a hydrocarbon oil processing system is shown, except that the hydrorefining reactor 13, the hydrorefining apparatus fractionating tower 21, and the apparatuses associated therewith are excluded;

the adopted hydrocarbon oil processing method comprises the following steps: referring to example 1, except that heavy oil feed 2 (properties as in table 1) was used instead of heavy oil feed 1; and the heavy oil raw material 2 is directly preheated to 260 ℃ without hydrotreating and enters the catalytic cracking unit as raw oil through a riser pipeline. The reaction conditions of the first, second and third risers in the hydrocarbon oil processing method are shown in table 3, and the product distribution after the device operates stably is shown in table 4.

Comparative example 1

The adopted hydrocarbon oil processing system comprises: referring to the system of fig. 3, the difference is that three riser reactors are not included, only one riser reactor is included;

the adopted hydrocarbon oil processing method comprises the following steps: the process of example 1 was used except that in the catalytic cracking unit, hydrotreated heavy wax oil was preheated to 260 ℃ as feed oil through the riser line into the catalytic cracking unit comprising only one riser reactor, and the mixture of LCO and LCN was refluxed into the riser reactor (hydrotreated heavy wax oil into the upper nozzle, mixture of LCO and LCN into the lower nozzle). The reaction conditions of the riser reactor in this system are shown in table 3; the product distribution after stable operation of the apparatus is shown in table 4.

Comparative examples 2 to 3

The adopted hydrocarbon oil processing system comprises: referring to the system of FIG. 3, the difference is that 1) the hydrofinishing reactor 13, the hydrofinishing unit fractionation column 21, and the equipment associated therewith are excluded; 2) three riser reactors are excluded, only one riser reactor is included;

the adopted hydrocarbon oil processing method comprises the following steps: the process of examples 2-3 was used except that in the catalytic cracking unit, heavy oil feedstock 2 was preheated to 260 ℃ as feed oil through the riser line into the catalytic cracking unit comprising only one riser reactor, and the mixture of LCO and LCN was refluxed into the riser reactor (hydrotreated heavy wax oil into the upper nozzle, mixture of LCO and LCN into the lower nozzle). The reaction conditions of the riser reactor in this system are shown in table 3; the product distribution after stable operation of the apparatus is shown in table 4.

Table 3.

Table 4.

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. A hydrocarbon oil processing method for producing ethylene and propylene in high yield comprises a catalytic cracking step and an oil-gas separation post-treatment step, and is characterized in that the catalytic cracking step comprises the following steps:

raw oil, recycled light diesel oil and a catalytic cracking catalyst are subjected to contact reaction in a first reaction zone to generate a first oil-gas mixture;

the recycled light naphtha, the raffinate oil extracted from heavy naphtha aromatic hydrocarbon and a catalytic cracking catalyst are subjected to contact reaction in a second reaction zone to generate a second oil-gas mixture; then the second oil-gas mixture and recycled C4 olefin are subjected to contact reaction in a third reaction zone to produce a third oil-gas mixture;

sending the first oil-gas mixture and the third oil-gas mixture into a settler for carrying out spent catalyst and oil-gas separation to obtain product oil gas;

in the oil-gas separation post-treatment step, the product oil gas is further divided into hydrogen, methane, products of ethylene and propylene, ethane, propane, C4 alkane, C4 olefin, light naphtha, heavy naphtha, light diesel oil, heavy diesel oil and the components, and the heavy naphtha is subjected to aromatic extraction to form light aromatic hydrocarbon and raffinate oil after heavy naphtha aromatic hydrocarbon extraction; wherein the C4 olefin and raffinate oil after extraction of light naphtha, light diesel oil and heavy naphtha aromatic hydrocarbon are refluxed to the corresponding reaction zone together.

2. The process according to claim 1, wherein an upper nozzle and a lower nozzle are included inside the first reaction zone, the raw oil is fed from the upper nozzle, and the recycled light diesel oil is fed from the lower nozzle.

3. The method of claim 2, wherein in the interval of the first reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is 0.2-5 seconds, and the agent-to-oil ratio is 10-100; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 0.8-9.8 seconds, and the agent-oil ratio is 4-50.

4. The method of claim 3, wherein in the interval of the first reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is 0.5-2 seconds, and the agent-oil ratio is 30-75; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 2-6 seconds; the agent-oil ratio is 5-12.

5. The process of claim 1 wherein the second reaction zone internally comprises an upper nozzle from which the light naphtha is fed and a lower nozzle from which the raffinate after the heavy naphtha aromatics extraction is fed.

6. The method of claim 5, wherein in the interval of the third reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is 0.2-5 seconds, and the agent-to-oil ratio is 10-100; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 0.8-9.8 seconds, and the agent-oil ratio is 5-60.

7. The process according to claim 6, wherein in the zone of the third reaction zone between the upper and lower nozzles, the residence time of the hydrocarbon oil is between 0.5 and 3 seconds; the agent-oil ratio is 30-70; in the interval between the upper nozzle and the outlet, the residence time of the hydrocarbon oil is 1-8 seconds, and the agent-oil ratio is 15-30.

8. The process of claim 1, wherein the residence time of the hydrocarbon oil in the third reaction zone is in the range of from 2 to 30 seconds.

9. The process of claim 8, wherein the residence time of the hydrocarbon oil in the third reaction zone is from 6 to 15 seconds.

10. The method of any one of claims 1 to 9,

the conditions of the contact reaction in the first reaction zone include: the reaction temperature is 500-750 ℃, and the pressure is 0.1-1.0 MPa;

the conditions of the contact reaction in the second reaction zone include: the reaction temperature is 500-750 ℃, and the pressure is 0.1-1.0 MPa;

the contact reaction conditions of the third reaction zone include: the reaction temperature is 490 ℃ and 740 ℃, and the pressure is 0.1-1.0 MPa.

11. The method of claim 10, wherein,

the conditions of the contact reaction in the first reaction zone include: the reaction temperature is 520-590 ℃, and the pressure is 0.2-0.4 MPa;

the conditions of the contact reaction in the second reaction zone include: the reaction temperature is 620-690 ℃, and the pressure is 0.2-0.4 MPa;

the contact reaction conditions of the third reaction zone include: the reaction temperature is 600-660 ℃, and the pressure is 0.2-0.4 MPa.

12. A process as claimed in any one of claims 1 to 9, wherein the catalytic cracking catalyst is a medium pore zeolite containing catalyst.

13. The process of any one of claims 1 to 9, wherein the catalytic cracking catalyst comprises, based on the total weight of the catalytic cracking catalyst, 2 to 60 wt% of zeolite, 10 to 99 wt% of inorganic oxide, 0 to 70 wt% of clay; wherein the zeolite comprises a medium pore zeolite and optionally a large pore zeolite.

14. The process as claimed in claim 13, wherein the zeolite contains 50 to 95 wt% of medium pore zeolite and 5 to 50 wt% of large pore zeolite, based on the total weight of the zeolite.

15. The process of claim 14 wherein the zeolite comprises from 70 to 95 weight percent of a medium pore zeolite and from 5 to 30 weight percent of a large pore zeolite.

16. The method of any one of claims 1 to 9, wherein the oil-gas separation post-treatment step further comprises contacting and reacting the ethane, propane and C4 alkanes with steam to produce ethylene and propylene as products and by-products of ethane, propane, C4 alkanes, C4 alkenes and steam cracked light naphtha.

17. The method as claimed in claim 16, wherein the temperature of the water vapor contact reaction is 700-1000 ℃.

18. The process of claim 16, wherein the reflux of ethane, propane, and C4 alkanes as by-products is further reacted with steam.

19. The process of claim 16 wherein the by-produced steam cracked light naphtha is returned to the second reaction zone and the by-produced C4 olefins are returned to the third reaction zone.

20. The process according to any one of claims 1 to 9, wherein in the step of aromatics extraction of the heavy naphtha to form light aromatics and raffinate oil after aromatics extraction of the heavy naphtha, the extraction solvent used is one or more selected from sulfolane, N-methylpyrrolidone, diethylene glycol ether, triethylene glycol ether, tetraethylene glycol, dimethyl sulfoxide and N-formyl morpholine ether.

21. The process of claim 20, wherein the operating conditions for aromatics extraction comprise: the temperature is 40-120 ℃, the volume ratio of the extraction solvent to the heavy naphtha is 0.5-5: 1.

22. the process according to any one of claims 1 to 9, wherein the feedstock oil is a hydrotreated heavy wax oil.

23. The process according to claim 22, wherein the feed oil is a hydrotreated heavy wax oil having a flowsheet greater than 350 ℃.

24. The method of claim 22, wherein the main operating conditions of the hydroprocessing comprise: the reaction temperature is 350 ℃ and 400 ℃, and the reaction is carried outThe pressure is 7-15MPa, and the volume space velocity is 0.5-1.5h^-1The hydrogen-oil ratio is 400-1500.

25. The method of claim 24, wherein the main operating conditions of the hydroprocessing comprise: the reaction temperature is 360-390 ℃, the reaction pressure is 9-13MPa, and the volume space velocity is 0.6-1.2h^-1The hydrogen-oil ratio is 600-1200.

26. The process of claim 22, wherein the hydrogenation catalyst employed in the hydroprocessing comprises: the hydrogenation catalyst comprises a carrier and oxides of metals in a VIB group and/or non-noble metals in a VIII group, wherein the oxides are loaded on the carrier, and the content of the metals in the VIB group in an oxidation state is 3-35 wt%, and the content of the non-noble metals in the VIII group in the oxidation state is 0.3-8 wt%, based on the total weight of the hydrogenation catalyst.

27. The method of claim 26, wherein the support is an alumina and/or silica-alumina support.

28. The method of claim 26 wherein the group vib metal is W.

29. The method of claim 26 wherein the group viii metal is Ni.

30. The process of claim 22, wherein the heavy diesel and the above components are refluxed to the step of hydrotreating.

31. A hydrocarbon oil processing system for the production of ethylene and propylene in high yield, said system comprising a catalytic cracking unit (30), characterized in that said catalytic cracking unit (30) comprises:

the first riser reactor (A) is provided with a raw oil inlet, a light diesel oil refining port and a first reaction mixture outlet;

the second riser reactor (B) is provided with a raffinate oil recycling port after heavy naphtha aromatic hydrocarbon is extracted, a light naphtha recycling port and a second reaction mixture outlet;

a third riser reactor (C) provided with a reaction mixture inlet connected with the second reaction mixture outlet, a C4 olefin recycle outlet, and a third reaction mixture outlet;

the settling device (32) is provided with a reaction mixture inlet, a spent agent outlet and a product oil gas outlet, wherein the reaction mixture inlet is respectively connected with the first reaction mixture outlet and the third reaction mixture outlet;

the regeneration device (33) is provided with a spent agent inlet connected with the spent agent outlet and a regeneration agent outlet, and the regeneration agent outlet is respectively connected with the catalyst inlets of the first riser reactor (A) and the second riser reactor (B);

the system also comprises an aromatic extraction unit which is used for extracting aromatic hydrocarbon from the heavy naphtha of the oil gas in the oil-gas separation post-treatment step and obtaining raffinate oil after the heavy naphtha aromatic hydrocarbon is extracted.

32. The system of claim 31 wherein the first riser reactor (a) is internally provided with an upper nozzle connected to the raw oil inlet and a lower nozzle connected to the light diesel oil recycle port.

33. The system of claim 31, wherein the second riser reactor (B) is internally provided with an upper nozzle connected to the light naphtha recycle port and a lower nozzle connected to the raffinate oil recycle port after the heavy naphtha aromatic extraction.

34. The system of any one of claims 32 to 33,

the system also comprises a raw oil supply unit, wherein the raw oil supply unit is connected with a raw oil inlet on the first riser reactor (A) in a matching way and is used for supplying raw oil into the first riser reactor (A);

the system further comprises a catalytic cracking fractionation unit (40), a gas separation unit, a light and heavy naphtha separation unit (50), and a light and heavy diesel separation unit (60);

the catalytic cracking fractionation unit (40) is connected with the settling device (32) in a matching manner and is used for fractionating the product oil gas separated by the settling device into dry gas (42a), liquefied gas (42b), naphtha (41b, 42c), diesel oil and the components (41 c);

the gas separation unit is connected with the catalytic cracking fractionation unit (40) in a matching way and is used for separating the dry gas (42a) and the liquefied gas (42b) into hydrogen, methane, products of ethylene and propylene, ethane, propane, C4 alkane and C4 alkene, wherein the C4 alkene flows back to the third riser reactor (C);

said light and heavy naphtha separation unit (50) cooperatively connected with said catalytic cracking fractionation unit (40) for separating naphtha (41B, 42c) into light naphtha (51a) and heavy naphtha (51B), wherein said light naphtha (51a) is refluxed to said second riser reactor (B);

the light and heavy diesel oil separation unit (60) is connected with the catalytic cracking fractionation unit (40) in a matching manner and is used for separating the diesel oil and the components (41c) into light diesel oil (61a) and heavy diesel oil and the components (61b), wherein the light diesel oil (61a) flows back to the first riser reactor (A).

35. The system of claim 34, further comprising a distillation cracking unit cooperatively connected with the gas separation unit for contacting ethane, propane and C4 alkanes with steam to react to produce ethylene and propylene as products and coproducing ethane, propane, C4 alkanes, C4 alkenes and steam cracked light naphtha, wherein the by-produced ethane, propane and C4 alkanes are refluxed to the distillation cracking unit; byproduct C4 olefins are refluxed to the third riser reactor (C) and byproduct steam cracked light naphtha is refluxed to the second riser reactor (B).

36. The system of claim 34, wherein said aromatics extraction unit is cooperatively associated with said light and heavy naphtha separation unit (50) for refluxing said heavy naphtha aromatics extracted raffinate to said second riser reactor (B).

37. The system according to claim 34, wherein the system further comprises a hydrofining unit (10) and a hydro-fractionation unit (20), the hydrofining unit (10) is cooperatively connected with the raw oil supply unit for hydrotreating the raw oil to obtain a hydrogenated product; the hydro-fractionation unit (20) is connected with the hydrofining unit (10) in a matching manner and is used for fractionating the hydrogenation product to obtain hydrotreated heavy wax oil; the hydrogenation fractionation unit (20) is connected with a raw oil inlet of a first riser reactor (A) in the catalytic cracking unit (30) in a matching manner, and is used for supplying the hydrotreating heavy wax oil to the first riser reactor (A).

38. The system according to claim 37, wherein the heavy diesel and the components thereof obtained by the light and heavy diesel separation unit (60) are returned to the hydrorefining unit.