阅读说明：本技术 劣质油的加工方法和系统 (Processing method and system for inferior oil ) 是由 陈学峰 魏晓丽 龚剑洪 侯焕娣 申海平 张执刚 张策 梁家林 戴立顺 张久顺 侯 于 2019-03-04 设计创作，主要内容包括：本发明涉及一种劣质油的加工方法和系统,该方法包括：将劣质油引入转化反应单元在临氢条件下进行转化反应,所得转化产物进行分离,得到重馏分；将所得重馏分进行萃取分离,得到改质油和残渣；将所得改质油和催化轻循环油引入加氢改质单元进行加氢改质,得到加氢改质油和加氢轻循环油；将所得加氢改质油和加氢轻循环油引入稀相输送床反应器中与催化裂化催化剂接触并进行催化裂化反应,得到待生催化剂和反应产物。采用本发明的方法和系统具有较高的高辛烷值汽油收率。(The invention relates to a processing method and a system of inferior oil, wherein the method comprises the following steps: introducing the inferior oil into a conversion reaction unit to carry out conversion reaction under the hydrogen condition, and separating the obtained conversion product to obtain heavy fraction; extracting and separating the obtained heavy fraction to obtain modified oil and residue; introducing the obtained modified oil and the catalytic light cycle oil into a hydro-modifying unit for hydro-modifying to obtain hydro-modified oil and hydro-light cycle oil; introducing the obtained hydrogenated modified oil and hydrogenated light cycle oil into a dilute phase transport bed reactor to contact with a catalytic cracking catalyst and carry out catalytic cracking reaction to obtain a spent catalyst and a reaction product. The method and the system have higher yield of the high-octane gasoline.)

1. A method of processing a poor quality oil, the method comprising:

introducing the inferior oil into a conversion reaction unit to carry out conversion reaction under the hydrogen condition, and separating the obtained conversion product to obtain heavy fraction with the distillation range of more than 350 ℃;

introducing the obtained heavy fraction into an extraction separation unit to contact with an extraction solvent and carrying out extraction separation to obtain modified oil and residue;

introducing the obtained modified oil and the catalytic light cycle oil into a hydro-modifying unit for hydro-modifying to obtain hydro-modified oil and hydro-light cycle oil;

introducing the obtained hydrogenated modified oil and hydrogenated light cycle oil into a dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and carry out catalytic cracking reaction to obtain a spent catalyst and a reaction product; according to the flow direction of the reaction materials, a feed inlet of the hydrogenated modified oil in the dilute phase transport bed reactor is positioned at the downstream of a feed inlet of the hydrogenated light cycle oil in the dilute phase transport bed reactor;

and returning the obtained spent catalyst to the dilute phase transport bed reactor after regeneration.

2. The process of claim 1, wherein said low grade oil comprises at least one selected from the group consisting of low grade crude oil, heavy oil, deoiled bitumen, coal derived oil, shale oil and petrochemical waste oil.

3. The process of claim 1 wherein the low quality oil meets one or more criteria selected from the group consisting of an API degree of less than 27, a distillation range of greater than 350 ℃, an asphaltene content of greater than 2 wt.%, and a heavy metal content of greater than 100 micrograms/gram based on the total weight of nickel and vanadium.

4. The process of claim 1, wherein said conversion reaction unit comprises a conversion reactor, said conversion reactor being a fluidized bed reactor;

the conversion reaction is carried out in the presence or absence of a conversion catalyst containing at least one selected from the group consisting of group VB metal compounds, group VIB metal compounds and group VIII metal compounds;

the conversion reaction conditions include: the temperature is 380-470 ℃, the hydrogen partial pressure is 10-25 MPa, and the volume space velocity of the inferior oil is 0.01-2 hours^-1The volume ratio of the hydrogen to the poor-quality oil is 500-5000, and the amount of the conversion catalyst is 10-50000 micrograms/g based on the weight of the poor-quality oil and calculated by the metal in the conversion catalyst.

5. The process of claim 1, wherein the conditions of the extractive separation comprise: the pressure is 3-12 MPa, the temperature is 55-300 ℃, and the extraction solvent is C₃-C₇The weight ratio of the hydrocarbon, the extraction solvent and the heavy fraction is (1-7): 1.

6. the process of claim 1, wherein the hydro-upgrading conditions comprise: hydrogen partial pressure of 5.0-20.0 MPa, reaction temperature of 330-450 deg.C, volume space velocity of 0.1-3 hr^-1The volume ratio of hydrogen to oil is 300-3000;

the catalyst used by the hydrogenation upgrading unit comprises a hydrofining catalyst and a hydrocracking catalyst, wherein the hydrofining catalyst comprises a carrier and an active metal component, and the active metal component is selected from VIB group metals and/or VIII group non-noble metals; the hydrocracking catalyst comprises zeolite, alumina, at least one group VIII metal component and at least one group VIB metal component.

7. The process of claim 6 wherein the hydrocracking catalyst comprises 3 to 60 wt% zeolite, 10 to 80 wt% alumina, 1 to 15 wt% nickel oxide and 5 to 40 wt% tungsten oxide based on the dry weight of the hydrocracking catalyst.

8. The process of claim 1, wherein the dilute phase transport bed reactor comprises, from bottom to top, a hydrogenated light cycle oil reaction zone and a hydrogenated upgraded oil reaction zone;

the conditions of the hydro-upgrading oil reaction zone comprise: the reaction temperature is 500-620 ℃, the reaction time is 0.1-5 seconds, the weight ratio of the catalyst to the oil is 5-15, and the weight ratio of the water to the oil is 0.05-0.3;

the conditions of the hydrogenation light cycle oil reaction zone comprise: the reaction temperature is 600-700 ℃, the reaction time is 1-10 seconds, the weight ratio of the catalyst to the oil is 1-50, and the weight ratio of the water to the oil is 0.01-0.5.

9. The process of claim 1 wherein the catalytic cracking catalyst comprises, on a dry weight basis and based on the weight of the catalytic cracking catalyst, from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay;

the zeolite comprises 50 to 100 wt% of a medium pore zeolite and 0 to 50 wt% of a large pore zeolite, based on the weight of the zeolite on a dry basis.

10. The process of claim 1, further comprising: returning the obtained residue to the conversion reaction unit for conversion reaction; and/or

And separating the obtained reaction product, and returning the separated catalytic light cycle oil to the hydro-upgrading unit for hydro-upgrading.

11. The process of claim 1, wherein the conversion of the conversion reaction is 30-70 wt%, the conversion being (weight of components with distillation range above 524 ℃ in the poor oil-weight of components with distillation range above 524 ℃ in the converted product)/weight of components with distillation range above 524 ℃ in the poor oil x 100 wt%; and/or

In the heavy fraction, the content of components with the distillation range between 350 ℃ and 524 ℃ is 20-60 wt%.

12. A processing system of inferior oil comprises a conversion reaction unit, an extraction separation unit, a hydrogenation modification unit and a dilute phase transport bed reactor;

the conversion reaction unit is provided with an inferior oil inlet, a hydrogen inlet, a light fraction outlet and a heavy fraction outlet, the extraction separation unit is provided with an extraction solvent inlet, a raw material inlet, a modified oil outlet and a residue outlet, the hydrogenation modification unit is provided with a hydrogen inlet, a modified oil inlet, a catalytic light cycle oil inlet, a hydrogenation modified oil outlet and a hydrogenation light cycle oil outlet, the dilute phase transport bed reactor is provided with a catalyst inlet, a first feed port, a second feed port and an oil agent outlet, and the first feed port of the dilute phase transport bed reactor is positioned at the upstream of the second feed port according to the flow direction of reaction materials;

the heavy fraction outlet of the conversion reaction unit is communicated with the raw material inlet of the extraction separation unit, the modified oil outlet of the extraction separation unit is communicated with the modified oil inlet of the hydrogenation modification unit, the first feed inlet of the dilute phase transport bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit, and the hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the second feed inlet of the dilute phase transport bed reactor.

Technical Field

The invention relates to a processing method and a system of inferior oil.

Background

In recent years, the quality of global gasoline for vehicles is rapidly improved, and the pace of upgrading the quality of oil products is obviously accelerated. At present, the quantity of catalytic cracking gasoline in China is about 70 percent of the total quantity of a gasoline pool, and the quality of the catalytic cracking gasoline plays a significant role in the overall level of the gasoline pool. The octane number RON of the catalytic cracking gasoline is 90-92 at most, the average octane number RON is 89-90, and the difference between the gasoline quality of the catalytic cracking gasoline and the gasoline quality of other developed countries is large, so that the octane number of the gasoline is improved, and the upgrading and updating of the gasoline are facilitated.

With the slow recovery of the world economy, the oil demand is slowly increased, and the supply and demand of the world oil market are basically kept loose. The international energy agency considers that, on the supply side, the crude oil production in non-european peck countries, represented by the united states, will continue to rise in the coming years, and the global crude oil demand will tend to tighten in 2022; on the demand side, the global crude oil demand will continuously rise in the next 5 years, and in 2019, 1 hundred million barrels per day will be broken through; the processing amount of unconventional oil and inferior heavy oil is increased year by year. Therefore, the method for producing the high-octane gasoline to the maximum extent by utilizing the unconventional oil or the inferior oil is the key and key point for expanding the production raw material source of the high-octane gasoline, guaranteeing the market demand and social supply, and improving the quality and the efficiency of petrochemical enterprises.

Chinese patent CN200910162163.9 discloses a method for processing inferior heavy oil by a combined process, which comprises subjecting inferior heavy oil raw material to solvent deasphalting to obtain deasphalted oil; preheating deasphalted oil, and then sequentially entering a first reaction zone and a second reaction zone of a catalytic conversion reactor to perform cracking reaction, hydrogen transfer reaction and isomerization reaction to obtain a catalytic wax oil product which accounts for 12-60% of the raw oil; and (3) hydrogenating the catalytic wax oil to obtain hydrogenated catalytic wax oil, introducing the hydrogenated catalytic wax oil into a catalytic conversion device, and further reacting to obtain a light fuel oil product. The method can produce propylene and high octane gasoline from poor raw oil to the maximum extent.

WO2015084779a1 discloses a process for the production of propylene and high octane gasoline, especially propylene, using a combination of solvent deasphalting and high severity catalytic cracking. The method comprises the following steps: mixing the vacuum residue and the solvent, and then performing solvent deasphalting treatment; the deasphalted oil rich in solvent enters a heavy oil deep catalytic cracking device for deep cracking reaction after the solvent is separated, and a target product rich in propylene and high-octane gasoline, particularly propylene, is obtained. The method comprises the steps of firstly carrying out solvent deasphalting treatment on residual oil, then realizing high-efficiency conversion of deasphalted oil and generating propylene and high-octane gasoline through a combined process, but not using and processing the deasphalted oil.

Chinese patent CN201410585111.3 discloses a method for treating heavy oil, which comprises subjecting a first heavy raw oil to fixed bed hydrogenation and fractionation to obtain fixed bed hydrogenation light oil and fixed bed hydrogenation residual oil; performing solvent deasphalting on the fixed bed hydrogenation residual oil to obtain deasphalted oil and deoiled asphalt; catalytically cracking the deasphalted oil and then fractionating to obtain products such as liquefied gas, gasoline and the like; hydrogenating and then fractionating hydrogenated raw oil containing deoiled asphalt, part or all of catalytic cracking heavy oil and second heavy raw oil; obtaining second hydrogenation light oil and second hydrogenation tail oil; at least part of the second hydrogenation tail oil returns to the catalytic cracking, and the gasoline yield can reach over 47 percent by the technical scheme.

In order to obtain more gasoline products from inferior oil, the prior art adopts a technical method combining solvent deasphalting and hydrotreating to provide a high-quality raw material for catalytic cracking, but the yield of deasphalted oil is low, the yield is limited from the economical point of the whole process, and in addition, the deasphalted oil is not well utilized, so that the utilization rate of the inferior oil in the prior art is low, and more residues are still generated. Therefore, there is a need to develop a green and efficient conversion technology for producing high-octane gasoline from low-grade oil, so as to increase the utilization rate of the low-grade oil and produce more high-value high-octane gasoline products.

Disclosure of Invention

The invention aims to provide a method and a system for processing inferior oil, and the method and the system have higher yield of high-octane gasoline.

In order to achieve the above object, the present invention provides a method for processing inferior oil, comprising:

introducing the inferior oil into a conversion reaction unit to carry out conversion reaction under the hydrogen condition, and separating the obtained conversion product to obtain heavy fraction with the distillation range of more than 350 ℃;

introducing the obtained heavy fraction into an extraction separation unit to contact with an extraction solvent and carrying out extraction separation to obtain modified oil and residue;

introducing the obtained modified oil and the catalytic light cycle oil into a hydro-modifying unit for hydro-modifying to obtain hydro-modified oil and hydro-light cycle oil;

introducing the obtained hydrogenated modified oil and hydrogenated light cycle oil into a dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and carry out catalytic cracking reaction to obtain a spent catalyst and a reaction product; according to the flow direction of the reaction materials, a feed inlet of the hydrogenated modified oil in the dilute phase transport bed reactor is positioned at the downstream of a feed inlet of the hydrogenated light cycle oil in the dilute phase transport bed reactor;

and returning the obtained spent catalyst to the dilute phase transport bed reactor after regeneration.

Optionally, the low-grade oil comprises at least one selected from the group consisting of low-grade crude oil, heavy oil, deoiled bitumen, coal-derived oil, shale oil, and petrochemical waste oil.

Optionally, the low quality oil meets one or more criteria selected from the group consisting of an API degree of less than 27, a distillation range of greater than 350 ℃, an asphaltene content of greater than 2 wt.%, and a heavy metal content of greater than 100 micrograms/gram, based on the total weight of nickel and vanadium.

Optionally, the conversion reaction unit comprises a conversion reactor, and the conversion reactor is a fluidized bed reactor;

the conversion reaction is carried out in the presence or absence of a conversion catalyst containing at least one selected from the group consisting of group VB metal compounds, group VIB metal compounds and group VIII metal compounds;

the conversion reaction conditions include: the temperature is 380-470 ℃, the hydrogen partial pressure is 10-25 MPa, and the volume space velocity of the inferior oil is 0.01-2 hours^-1The volume ratio of the hydrogen to the poor-quality oil is 500-5000, and the amount of the conversion catalyst is 10-50000 micrograms/g based on the weight of the poor-quality oil and calculated by the metal in the conversion catalyst.

Optionally, the extraction separation conditions include: the pressure is 3-12 MPa, the temperature is 55-300 ℃, and the extraction solvent is C₃-C₇The weight ratio of the hydrocarbon, the extraction solvent and the heavy fraction is (1-7): 1.

optionally, the hydro-upgrading conditions include: the hydrogen partial pressure is 5.0-20.0 MPa, the reaction temperature is 330-^-1The volume ratio of hydrogen to oil is 300-3000;

the catalyst used by the hydrogenation upgrading unit comprises a hydrofining catalyst and a hydrocracking catalyst, wherein the hydrofining catalyst comprises a carrier and an active metal component, and the active metal component is selected from VIB group metals and/or VIII group non-noble metals;

the hydrocracking catalyst comprises zeolite, alumina, at least one group VIII metal component and at least one group VIB metal component;

optionally, the hydrocracking catalyst comprises 3-60 wt% zeolite, 10-80 wt% alumina, 1-15 wt% nickel oxide and 5-40 wt% tungsten oxide, based on the dry weight of the hydrocracking catalyst.

In the hydrogenation modification catalyst, the filling volume ratio of the hydrofining catalyst to the hydrocracking catalyst is 1-5:1, according to the flow direction of the reaction materials, the hydrofining catalyst is filled at the upstream of the hydrocracking catalyst.

Optionally, the dilute-phase transport bed reactor comprises a hydrogenation light cycle oil reaction zone and a hydrogenation modified oil reaction zone from bottom to top;

the conditions of the hydro-upgrading oil reaction zone comprise: the reaction temperature is 500-620 ℃, the reaction time is 0.1-5 seconds, the weight ratio of the catalyst to the oil is 5-15, and the weight ratio of the water to the oil is 0.05-0.3;

the conditions of the hydrogenation light cycle oil reaction zone comprise: the reaction temperature is 600-700 ℃, the reaction time is 1-10 seconds, the weight ratio of the catalyst to the oil is 1-50, and the weight ratio of the water to the oil is 0.01-0.5.

Optionally, the catalytic cracking catalyst comprises, on a dry basis weight basis and based on the weight of the catalytic cracking catalyst, from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay;

the zeolite comprises 50 to 100 wt% of a medium pore zeolite and 0 to 50 wt% of a large pore zeolite, based on the weight of the zeolite on a dry basis.

Optionally, the method further includes: returning the obtained residue to the conversion reaction unit for conversion reaction; and/or

And separating the obtained reaction product, and returning the separated catalytic light cycle oil to the hydro-upgrading unit for hydro-upgrading.

Optionally, the conversion rate of the conversion reaction is 30-70 wt%, and the conversion rate of the conversion reaction is (the weight of the component with the distillation range above 524 ℃ in the inferior oil-the weight of the component with the distillation range above 524 ℃ in the conversion product)/the weight of the component with the distillation range above 524 ℃ in the inferior oil x 100 wt%; and/or

In the heavy fraction, the content of components with the distillation range between 350 ℃ and 524 ℃ is 20-60 wt%.

The invention also provides a processing system of the inferior oil, which comprises a conversion reaction unit, an extraction separation unit, a hydrogenation modification unit and a dilute phase transport bed reactor;

the conversion reaction unit is provided with an inferior oil inlet, a hydrogen inlet, a light fraction outlet and a heavy fraction outlet, the extraction separation unit is provided with an extraction solvent inlet, a raw material inlet, a modified oil outlet and a residue outlet, the hydrogenation modification unit is provided with a hydrogen inlet, a modified oil inlet, a catalytic light cycle oil inlet, a hydrogenation modified oil outlet and a hydrogenation light cycle oil outlet, the dilute phase transport bed reactor is provided with a catalyst inlet, a first feed port, a second feed port and an oil agent outlet, and the first feed port of the dilute phase transport bed reactor is positioned at the upstream of the second feed port according to the flow direction of reaction materials;

the heavy fraction outlet of the conversion reaction unit is communicated with the raw material inlet of the extraction separation unit, the modified oil outlet of the extraction separation unit is communicated with the modified oil inlet of the hydrogenation modification unit, the first feed inlet of the dilute phase transport bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit, and the hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the second feed inlet of the dilute phase transport bed reactor.

Compared with the prior art, the invention has the following advantages:

1. the method can process the inferior oil with high metal and high asphaltene content, realizes the high-efficiency lightening of the asphaltene by introducing the pre-conversion reaction of the inferior oil, and greatly reduces the residue amount.

2. The distillation range and the composition of the raw materials for extraction and separation are optimized, the extraction and separation process is easy to operate, and the physical properties of the residues obtained in the extraction and separation process are improved, so that the subsequent division of labor is facilitated.

3. Can provide high-quality raw materials without metal and asphaltene for catalytic cracking, and realizes the maximized production of high-octane gasoline products with high added value.

4. The light cycle oil can be used as a catalytic cracking feed after being subjected to hydro-upgrading to further produce high-octane gasoline, so that the pressure of surplus diesel markets and insufficient supply of the high-octane gasoline is relieved.

5. Not only can realize the efficient green conversion of the low-quality oil, but also can realize the production of high value-added products, namely high-octane gasoline from the low-quality oil.

Additional features and advantages of the invention will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the invention and are incorporated in and constitute a part of this specification, illustrate embodiments of the invention and together with the description serve to explain the principles of the invention and not to limit the invention. In the drawings:

FIG. 1 is a schematic flow diagram including an embodiment of the method of the present invention and further including a schematic structural diagram of an embodiment of the system of the present invention.

Description of the reference numerals

1 dilute phase transport bed reactor 2 regenerator 3 settler

4 stripping section 5 degassing tank 6 cyclone separator

7 gas collection chamber 8 spent inclined tube 9 spent slide valve

10 line 11 line 12 regenerative chute

13 regenerative slide valve 14 line 15 line

16 line 17 line 18 line

19 line 20 large oil and gas line 21 line

22 air distributor 23 line 24 cyclone

25 flue gas pipeline

30 conversion reaction unit 31 line 32 line

33 line 34 line 35 line

36 extractive separation unit 37 line 38 line

39 line 40 hydro-upgrading unit 41 line

Detailed Description

The following describes in detail specific embodiments of the present invention. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.

The invention provides a processing method of inferior oil, which comprises the following steps:

introducing the inferior oil into a conversion reaction unit to carry out conversion reaction under the hydrogen condition, and separating the obtained conversion product to obtain heavy fraction with the distillation range of more than 350 ℃;

introducing the obtained heavy fraction into an extraction separation unit to contact with an extraction solvent and carrying out extraction separation to obtain modified oil and residue;

introducing the obtained modified oil and the catalytic light cycle oil into a hydro-modifying unit for hydro-modifying to obtain hydro-modified oil and hydro-light cycle oil;

introducing the obtained hydrogenated modified oil and hydrogenated light cycle oil into a dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and carry out catalytic cracking reaction to obtain a spent catalyst and a reaction product; according to the flow direction of the reaction materials, a feed inlet of the hydrogenated modified oil in the dilute phase transport bed reactor is positioned at the downstream of a feed inlet of the hydrogenated light cycle oil in the dilute phase transport bed reactor;

and returning the obtained spent catalyst to the dilute phase transport bed reactor after regeneration.

According to the present invention, low-grade oils are well known to those skilled in the art, and for example, the low-grade oils may include at least one selected from the group consisting of low-grade crude oil, heavy oil, deoiled bitumen, coal-derived oil, shale oil, and petrochemical waste oil. The heavy oil refers to distillate oil or residual oil with a boiling point above 350 ℃, and the distillate oil generally refers to fraction products obtained by atmospheric distillation and vacuum distillation of crude oil or secondary processing oil, such as heavy diesel oil, heavy gas oil, lubricating oil fraction or cracking raw materials and the like; residuum refers to the bottoms from atmospheric and vacuum distillation of crude oil, and atmospheric distillation bottoms are generally referred to as long residues (typically large boiling)A fraction at 350 ℃), which is generally referred to as a vacuum residue (generally, a fraction having a boiling point of more than 500 ℃ or 524 ℃), and which may be at least one selected from a topped crude oil, a heavy oil derived from oil sand bitumen, and a heavy oil having a primary boiling point of more than 350 ℃, the topped crude oil being an oil discharged from the bottom of a primary tower or the bottom of a flash tower when the crude oil is fractionated in an atmospheric and vacuum distillation process; the inferior crude oil is thick oil, which refers to crude oil with high content of asphaltene and colloid and high viscosity, and the density of the ground is generally more than 0.943 g/cm at 20 DEG C³The crude oil with the viscosity of the underground crude oil being more than 50 centipoises is called thick oil; the deasphalted oil is rich in asphaltene and rich in aromatic components, which are obtained by contacting, dissolving and separating raw oil with a solvent and at the bottom of an extraction tower in a solvent deasphalting device, and can be divided into propane deasphalted asphalt, butane deoiled asphalt, pentane deoiled asphalt and the like according to the difference of the types of the solvents; the coal derived oil is a liquid fuel obtained by taking coal as a raw material and performing chemical processing, and can be at least one selected from coal liquefied oil generated by coal liquefaction and coal tar generated by coal pyrolysis; shale oil is brown sticky paste obtained by low-temperature dry distillation of oil shale, and has pungent odor and high nitrogen content; the petrochemical waste oil may be at least one selected from the group consisting of petrochemical waste oil sludge, petrochemical oil residue, and refined products thereof. Other low quality oils known to those skilled in the art may also be used alone or in combination as low quality oils for the conversion reaction and the present invention is not further described herein. The low quality oil preferably meets one or more criteria selected from the group consisting of an API degree of less than 27, a distillation range of greater than 350 ℃ (preferably greater than 500 ℃, more preferably greater than 524 ℃), an asphaltene content of greater than 2 wt.% (preferably greater than 5 wt.%, more preferably greater than 10 wt.%, even more preferably greater than 15 wt.%), and a heavy metal content of greater than 100 micrograms/gram, based on the total weight of nickel and vanadium.

According to the invention, the conversion reaction is essentially a thermal conversion reaction, which means that the poor oil is thermally converted in the presence of hydrogen and conversion products are obtained which contain at least heavy fractions and which may also contain light fractions having a lower distillation range than the heavy fractions.

In one embodiment, the conversion reaction unit comprises a conversion reactor, and the conversion reaction can be carried out in a fluidized bed reactor by using a solid-liquid suspension as a catalyst, so that the conversion reactor can be a fluidized bed reactor, and the fluidized bed reactor is a reactor in which reaction raw materials and the catalyst are reacted in a flowing state, and generally comprises a slurry bed reactor, a suspended bed reactor and a fluidized bed reactor, and the slurry bed reactor is preferred in the invention.

In one embodiment, the conversion reaction is carried out in the presence or absence of a conversion catalyst comprising at least one member selected from the group consisting of a group VB metal compound, a group VIB metal compound, and a group VIII metal compound; preferably at least one of Mo compound, W compound, Ni compound, Co compound, Fe compound, V compound and Cr compound, for example at least one selected from molybdenum naphthenate, nickel naphthenate, ammonium molybdate, organic molybdenum, organic vanadium and hematite; the conditions of the conversion reaction may include: the temperature is 380-470 ℃, preferably 400-440 ℃, the hydrogen partial pressure is 10-25 MPa, preferably 13-20 MPa, and the volume space velocity of the inferior oil is 0.01-2 hours^-1Preferably 0.1 to 1.0 hour^-1The volume ratio of the hydrogen to the poor-quality oil is 500-5000, preferably 800-2000, and the amount of the conversion catalyst is 10-50000 micrograms/g, preferably 30-25000 micrograms/g based on the weight of the poor-quality oil and the metal in the conversion catalyst.

In one embodiment, when the process of the present invention is carried out in a conversion reaction unit, it is generally carried out as follows: the poor oil, hydrogen and the conversion catalyst enter a conversion reactor of a conversion reaction unit for reaction, and the reaction product is separated into a gas product and a liquid product to finally obtain heavy fraction with the distillation range of more than 350 ℃.

According to the invention, the extractive separation, which can be carried out in an extraction solvent at a certain temperature and under a certain pressure, preferably in countercurrent contact extraction of the extraction solvent with the heavy fraction, can be carried out in any extraction apparatus, for example in any extraction apparatus, is used for separating easily processable upgrading oil from the heavy fraction and for carrying out the conversion reaction by external throwing or returning the residueThe extraction separation is carried out in an extraction tower, and the conditions of the extraction separation comprise: the pressure is 3-12 MPa, preferably 3.5-10 MPa, the temperature is 55-300 deg.C, preferably 70-220 deg.C, and the extraction solvent is C₃-C₇A hydrocarbon, preferably C₃-C₅Alkane and C₃-C₅At least one of olefins, more preferably C₃-C₄Alkane and C₃-C₄At least one olefin, the weight ratio of the extraction solvent to the heavy fraction being (1-7): 1, preferably (1.5-5): 1. other conventional extraction methods can be adopted by the person skilled in the art for extraction, and the description of the invention is omitted.

In one embodiment, when the process of the present invention is carried out in an extractive separation unit, the heavy fraction having a distillation range of greater than 350 ℃ obtained from the conversion reaction unit is sent to the extractive separation unit for countercurrent contact with an extraction solvent for extractive separation to obtain an upgraded oil and a residue. The residue can be returned to the conversion reaction unit for further conversion reaction, and the modified oil is sent to the hydro-upgrading unit.

The catalytic light cycle oil is well known to those skilled in the art, the source of the catalytic light cycle oil is not particularly limited in the present invention, and in one embodiment, the obtained reaction product is separated, and the separated catalytic light cycle oil is returned to the hydro-upgrading unit and the upgraded oil is subjected to hydro-upgrading.

According to the present invention, the hydro-upgrading can be carried out in any manner known in the art, without particular limitation, and can be carried out in any hydrogenation apparatus known in the art (e.g., fixed bed reactor, fluidized bed reactor), which can be reasonably selected by one skilled in the art. The modified oil and the catalytic light cycle oil are subjected to hydro-modification in the hydro-modification unit, and the hydro-modification conditions and the used catalyst of the modified oil and the catalytic light cycle oil can be the same or different, and can be carried out in the same hydro-modification device or different hydro-modification devices. The hydro-upgrading conditions include: the hydrogen partial pressure is 5.0-20.0 MPa, the preferential pressure is 8-15 MPa, the reaction temperature is 330-450 ℃, the preferential pressure is 350-420 ℃, and the volume space velocity is 0.1-3 hours^-1Preferably 0.3 to 1.5 hours^-1Volume ratio of hydrogen to oil300-3000, preferably 800-1500; the catalyst used in the hydro-upgrading unit may include a hydrofinishing catalyst and a hydrocracking catalyst, and for example, any catalyst conventionally used in the art for this purpose may be used or may be manufactured according to any manufacturing method conventionally known in the art, and the amounts of the hydrofinishing catalyst and the hydrocracking catalyst used in the step may be referred to the conventional knowledge in the art and are not particularly limited. Specifically, the hydrorefining catalyst may comprise a carrier and an active metal component, and more specifically, examples of the active metal component include metals of group VIB and non-noble metals of group VIII of the periodic table, and particularly, a combination of nickel and tungsten, a combination of nickel, tungsten and cobalt, a combination of nickel and molybdenum, or a combination of cobalt and molybdenum. These active metal components may be used singly or in combination in any ratio. Examples of the carrier include alumina, silica, and amorphous silica-alumina. These carriers may be used singly or in combination in any ratio. The respective contents of the carrier and the active metal component are not particularly limited in the present invention, and conventional knowledge in the art can be referred to. The hydrocracking catalyst generally contains a cracking functional component, such as a zeolite, and specifically, the hydrocracking catalyst may include a zeolite, alumina, at least one group VIII metal component and at least one group VIB metal component, and the zeolite may be a Y-type zeolite. The hydrocracking catalyst may comprise 3-60 wt% zeolite, 10-80 wt% alumina, 1-15 wt% nickel oxide and 5-40 wt% tungsten oxide based on the dry weight of the hydrocracking catalyst, although other compositions of hydrocracking catalysts may be employed by those skilled in the art. The invention has no special requirements on the loading sequence and the loading proportion of the hydrofining catalyst and the hydrocracking catalyst, for example, the hydrofining catalyst and the hydrocracking catalyst can be loaded in any ratio in a mixing way or loaded upstream and downstream, and the hydrofining catalyst is preferably loaded upstream of the hydrocracking catalyst, for example: in the catalyst used in the hydrogenation modification unit, the filling volume ratio of the hydrofining catalyst to the hydrocracking catalystThe ratio of the hydrogen refining catalyst to the hydrocracking catalyst is 1-5:1, and the hydrogen refining catalyst is filled at the upstream of the hydrocracking catalyst according to the flow direction of a reaction material.

In one embodiment, when the process of the present invention is carried out in a hydro-upgrading unit, the upgraded oil from the extractive separation unit is reacted with a hydro-upgrading catalyst to obtain a hydro-upgraded oil and a hydro-light cycle oil, which are sent to different reaction zones of the catalytic cracking reactor.

According to the present invention, the dilute phase transport bed reactor is a catalytic cracking reactor well known to those skilled in the art, and includes, for example, from bottom to top, a hydrogenated light cycle oil reaction zone and a hydrogenated upgraded oil reaction zone, both of which are bounded by a feed inlet of the hydrogenated upgraded oil in the dilute phase transport bed reactor; the conditions of the catalytic cracking reaction may be selected as desired, for example, the conditions of the hydro-upgrading oil reaction zone may include: the reaction temperature is 500-620 ℃, the reaction pressure is 0.2-1.2 MPa, the reaction time is 0.1-5 seconds, the catalyst-oil weight ratio (the weight ratio of the catalytic cracking catalyst to the hydro-upgrading oil) is 5-15, and the water-oil weight ratio (the weight ratio of the steam to the hydro-upgrading oil) is 0.05-0.3; the conditions of the hydrogenated light cycle oil reaction zone may include: the reaction temperature is 600-700 ℃, the reaction pressure is 0.2-1.2 MPa, the reaction time is 1-10 seconds, the catalyst-oil weight ratio (the weight of the catalytic cracking catalyst to the weight of the hydrogenated light cycle oil) is 1-50, and the water-oil weight ratio (the weight of the water vapor to the weight of the hydrogenated light cycle oil) is 0.01-0.5.

In one embodiment, when the process of the present invention is carried out in a catalytic cracking reactor, it is generally carried out as follows: the catalytic cracking catalyst enters a pre-lifting section of a catalytic cracking reactor, flows upwards under the action of a pre-lifting medium, preheated hydrogenated light cycle oil and atomized steam are injected into the bottom of the reactor together, the preheated hydrogenated light cycle oil and the atomized steam are firstly contacted with a regenerated catalyst to carry out catalytic cracking reaction and simultaneously flow upwards, then are contacted with preheated hydrogenated modified oil to carry out catalytic cracking reaction, the reacted material flow enters a cyclone separator through the outlet of the reactor, the separated reaction product is led out of a device to be further separated to obtain a gasoline product with high octane number, light cycle oil and other products, and the light cycle oil is sent to a hydrogenation modification unit to carry out hydrogenation modification; the separated spent catalyst enters a regenerator for coke burning regeneration, and the regenerated catalyst with recovered activity returns to the catalytic cracking reactor for recycling.

According to the invention, the catalytic cracking reactor may also be injected with steam, preferably in the form of atomized steam. The weight ratio of injected steam to hydroupgraded oil (hydrogenated light cycle oil) may be from 0.01 to 1, preferably from 0.05 to 0.5.

In the present invention, the spent catalyst and the reaction product are generally separated to obtain the spent catalyst and the reaction product, and then the obtained reaction product is separated by a subsequent separation system to obtain a high octane gasoline product, a light cycle oil and other products.

According to the invention, the method of the invention also preferably comprises: regenerating the spent catalyst; and preferably at least a part of the catalytic cracking catalyst is regenerated, for example, the whole of the catalyst may be regenerated catalyst.

According to the invention, the method of the invention also preferably comprises: the regenerated catalyst obtained by regeneration is stripped (generally by steam stripping) to remove impurities such as gas.

According to the invention, in the regeneration process, oxygen-containing gas is generally introduced from the bottom of the regenerator, the oxygen-containing gas can be air, for example, after being introduced into the regenerator, the catalyst to be generated is contacted with oxygen for scorching and regeneration, the gas-solid separation is carried out on the upper part of the regenerator on the flue gas generated after the scorching and regeneration of the catalyst, and the flue gas enters a subsequent energy recovery system.

According to the invention, the regeneration operation conditions of the spent catalyst can be as follows: the regeneration temperature is 550-750 ℃, preferably 600-730 ℃, and more preferably 650-700 ℃; the gas superficial linear velocity is 0.5 to 3 m/s, preferably 0.8 to 2.5 m/s, more preferably 1 to 2 m/s, and the average residence time of the spent catalyst is 0.6 to 3 minutes, preferably 0.8 to 2.5 minutes, more preferably 1 to 2 minutes.

Catalytic cracking catalysts are well known to those skilled in the art in light of the present invention, and conventional catalysts may be employed in the practice of the present inventionAlternatively, for example, the catalytic cracking catalyst may comprise, on a dry weight basis and based on the weight of the catalytic cracking catalyst, from 1 to 60 weight percent zeolite, from 5 to 99 weight percent inorganic oxide, and from 0 to 70 weight percent clay; the zeolite is taken as an active component, preferably the zeolite is selected from medium-pore zeolite and/or large-pore zeolite, and the medium-pore zeolite accounts for 50-100 wt% of the total weight of the zeolite, preferably the medium-pore zeolite accounts for 70-100 wt% of the total weight of the zeolite, and the large-pore zeolite accounts for 0-50 wt% of the total weight of the zeolite, preferably the large-pore zeolite accounts for 0-30 wt% of the total weight of the zeolite. The medium pore size zeolite and the large pore size zeolite are defined by the convention in the art, i.e., the medium pore size zeolite has an average pore size of 0.5 to 0.6 nm and the large pore size zeolite has an average pore size of 0.7 to 1.0 nm. For example, the large pore zeolite may be selected from a mixture of one or more of the group of zeolites consisting of rare earth Y (rey), rare earth hydrogen Y (rehy), ultrastable Y obtained by different methods, high silicon Y. The intermediate pore size zeolite may be selected from zeolites having the MFI structure, such as ZSM-series zeolites and/or ZRP zeolites, which may also be modified with non-metallic elements such as phosphorus and/or transition metal elements such as iron, cobalt, nickel, as described in more detail in connection with ZRP, see U.S. Pat. No. 5,232,675, and the ZSM-series zeolites may be selected from one or more mixtures of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites of similar structure, as described in more detail in connection with ZSM-5, see U.S. Pat. No. 3,702,886. In the present invention, the inorganic oxide is preferably selected from silicon dioxide (SiO) as a binder₂) And/or aluminum oxide (Al)₂O₃). In the present invention, the clay is preferably selected from kaolin and/or halloysite as a matrix (i.e., carrier).

According to the invention, the conversion rate of the conversion reaction can be 30-70 wt%, the conversion rate of the conversion reaction is (weight of component with distillation range above 524 ℃ in inferior oil-weight of component with distillation range above 524 ℃ in conversion product)/weight of component with distillation range above 524 ℃ in inferior oil x 100 wt%; and/or the fraction of the heavy fraction may have a fraction of between 20 and 60 wt.% of components having a distillation range between 350 and 524 ℃. The invention can maintain the long-time operation of the system under the condition of reducing the outward throwing of the residue as much as possible and improving the resource utilization rate, the conversion reactor and the extraction separation unit are the key for determining whether the system can operate for a long time, the conversion rate of the conversion reactor is as high as possible under the condition of system stability, light fractions less than 350 ℃ in the first separation product entering the extraction separation unit are not too much, otherwise, the light fractions pollute the solvent, black oil is generated in the extraction separation process, heavy fractions with the distillation range of 350 ℃ and 524 ℃ are more, otherwise, the residue is not easy to flow and the conversion reaction is not easy to be carried out in the conversion reactor. The conversion rate of the conversion reaction is high, coke is easy to form, and the system running time is reduced, and the conversion rate is low, the outward throwing residue is easy to be too much, and the unit time modification efficiency is reduced.

The invention also provides a processing system of the inferior oil, which comprises a conversion reaction unit, an extraction separation unit, a hydrogenation modification unit and a dilute phase transport bed reactor;

the conversion reaction unit is provided with an inferior oil inlet, a hydrogen inlet, a light fraction outlet and a heavy fraction outlet, the extraction separation unit is provided with an extraction solvent inlet, a raw material inlet, a modified oil outlet and a residue outlet, the hydrogenation modification unit is provided with a hydrogen inlet, a modified oil inlet, a catalytic light cycle oil inlet, a hydrogenation modified oil outlet and a hydrogenation light cycle oil outlet, the dilute phase transport bed reactor is provided with a catalyst inlet, a first feed port, a second feed port and an oil agent outlet, and the first feed port of the dilute phase transport bed reactor is positioned at the upstream of the second feed port according to the flow direction of reaction materials;

the heavy fraction outlet of the conversion reaction unit is communicated with the raw material inlet of the extraction separation unit, the modified oil outlet of the extraction separation unit is communicated with the modified oil inlet of the hydrogenation modification unit, the first feed inlet of the dilute phase transport bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit, and the hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the second feed inlet of the dilute phase transport bed reactor.

According to the invention, the equipment used for catalytic cracking can also comprise parts such as oil separating equipment and a stripper, the oil separating equipment can be a settler with a cyclone separator, and other conventional parts can be arranged as required by a person skilled in the art, and the details are not repeated.

The following detailed description of embodiments of the invention refers to the accompanying drawings. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.

As shown in fig. 1, the poor oil is transferred to the shift reaction unit 30 through a line 31, a shift catalyst through a line 32, and hydrogen through a line 33 for shift reaction. The light fraction separated from the reaction product is led out through a pipeline 34, and the heavy fraction separated with the distillation range of more than 350 ℃ is conveyed to an extraction separation unit 36 through a pipeline 35 to be in countercurrent contact with an extraction solvent from a pipeline 37 for extraction separation, so that the modified oil and the residue are obtained. The residue is recycled via line 38 to the shift reaction unit 30 for continued shift reaction with the low grade oil. The modified oil enters a hydrogenation modification unit 40 for hydrogenation modification through a pipeline 39 and the catalytic light cycle oil through a pipeline 26, the hydrogenated light cycle oil is sent into the dilute phase transport bed reactor 1 through a pipeline 16 and the hydrogenated modified oil through a pipeline 18, and other products are led out through a pipeline 41.

The pre-lifting medium enters the bottom of the dilute phase transport bed reactor 1 through a pipeline 14, the regenerated catalyst from a pipeline 12 enters the dilute phase transport bed reactor 1 after being regulated by a regeneration slide valve 13 and then moves upwards in an accelerated manner along the dilute phase transport bed reactor under the lifting action of the pre-lifting medium, the preheated hydrogenated light cycle oil is injected into the bottom of the dilute phase transport bed reactor 1 through a pipeline 16 together with the atomized steam from a pipeline 15, is mixed with the existing material flow of the dilute phase transport bed reactor 1, and is subjected to cracking reaction on the hot catalyst and moves upwards in an accelerated manner, and then is contacted with the preheated hydrogenated modified oil from a pipeline 18 and the atomized steam from a pipeline 17 again to perform catalytic cracking reaction; the generated reaction product and the inactivated spent catalyst enter a cyclone separator 6 in a settler 3 to realize the separation of the spent catalyst and the reaction product, the reaction product enters an air collection chamber 7, and the fine powder of the catalyst returns to the settler. Spent catalyst in the settler flows to the stripping section 4 where it is contacted with steam from line 19. The reaction product stripped from the spent catalyst enters the gas collection chamber 7 after passing through the cyclone separator. The stripped spent catalyst enters the regenerator 2 through a spent inclined tube 8 after being regulated by a spent slide valve 9, air from a pipeline 21 enters the regenerator 2 after being distributed by an air distributor 22, coke on the spent catalyst in a dense bed layer at the bottom of the regenerator 2 is burned out, the deactivated spent catalyst is regenerated, and flue gas enters a subsequent energy recovery system through an upper gas flue gas pipeline 25 of a cyclone separator 24. Wherein the pre-lifting medium may be dry gas, water vapor or a mixture thereof.

The regenerated catalyst enters a degassing tank 5 through a pipeline 10 communicated with a catalyst outlet of a regenerator 2, and is contacted with a stripping medium from a pipeline 23 at the bottom of the degassing tank 5 to remove flue gas carried by the regenerated catalyst, the degassed regenerated catalyst circulates to the bottom of a dilute phase transport bed reactor 1 through a pipeline 12, the catalyst circulation amount can be controlled through a regeneration slide valve 13, gas in the degassing tank 5 returns to the regenerator 2 through a pipeline 11, reaction products in a gas collection chamber 7 enter a subsequent separation system through a large oil-gas pipeline 20, and light cycle oil obtained by separation enters a hydro-upgrading unit through a pipeline 26.

The following examples further illustrate the process but do not limit the invention.