CN111647430B - Treatment method and system for inferior oil - Google Patents

Abstract

The invention relates to a method and a system for treating 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 into the bottom of a main dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction; and introducing the obtained hydrogenated light cycle oil into a secondary dilute phase transport bed reactor to contact with a catalytic cracking catalyst and carrying out catalytic cracking reaction. The method and the system have higher propylene yield and high octane number gasoline yield.

Description

Treatment method and system for inferior oil

Technical Field

The invention relates to a method and a system for treating 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.

Meanwhile, ethylene and propylene are important petrochemical industry base stocks. Currently, about 98% of the world's ethylene comes from tubular furnace steam cracking technology, with 46% naphtha and 34% ethane in the ethylene production feed. About 62% of the propylene comes from the co-production of ethylene by steam cracking. The steam cracking technology is perfected day by day, and is a process of consuming a large amount of energy, and is limited by the use of high-temperature materials, and the potential of further improvement is very small. In addition, the fixed ethylene/propylene ratio of steam cracking is also difficult to accommodate and meet the faster growing propylene demand.

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 propylene and high-octane gasoline to the maximum extent by utilizing unconventional oil or poor oil is the key and key point for broadening production raw material sources of propylene and high-octane gasoline for petrochemical enterprises, guaranteeing market demand and social supply, and improving quality and efficiency.

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 provided by the patent furthest produces propylene and high-octane gasoline from poor-quality raw oil.

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 patent method firstly carries out solvent deasphalting treatment on residual oil, and then realizes the high-efficiency conversion of deasphalted oil and the generation of propylene and high-octane gasoline through a combined process, but the deasphalted oil is not used and processed.

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 high-value propylene and high-octane 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 viewpoint of economy of the whole process, in addition, the deasphalted oil is not well utilized, so that the utilization rate of the inferior oil in the prior art is not high, and more residues are still generated. Therefore, there is a need to develop a green and efficient conversion technology for producing propylene and high-octane gasoline from low-grade oil, so as to increase the utilization rate of the low-grade oil and produce more high-value propylene and high-octane gasoline products.

Disclosure of Invention

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

In order to achieve the above object, the present invention provides a method for treating 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 into the bottom of a main dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a first catalyst to be generated and a first reaction product;

introducing the obtained hydrogenated light cycle oil into a secondary dilute phase transport bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a second spent catalyst and a second reaction product;

and regenerating the first catalyst to be regenerated and the second catalyst to be regenerated, and returning the regenerated first catalyst to the main dilute phase conveying bed reactor and the auxiliary dilute phase conveying bed reactor.

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 hydrogenation modification catalyst comprises a hydrogenation refining catalyst and a hydrocracking catalyst, wherein the hydrogenation refining 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;

in the hydrogenation modification catalyst, the filling volume ratio of the hydrogenation refining 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 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.

Optionally, the conditions of the catalytic cracking reaction include: the reaction temperature is 500-650 ℃, the reaction time is 1-8 seconds, and the weight ratio of the catalytic cracking catalyst to the hydro-upgrading oil is 1-50;

the conditions of the catalytic cracking reaction include: the reaction temperature is 560 ℃ and 750 ℃, the reaction time is 1-10 seconds, and the weight ratio of the catalytic cracking catalyst to the hydrogenation light cycle oil is 1-100;

the feed weight ratio of the hydrogenated modified oil to the hydrogenated light cycle oil is 1: (0.1-0.5).

Optionally, the catalytic cracking catalyst and the catalytic cracking catalyst each independently comprise, on a dry basis weight 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 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 first reaction product and/or second 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 treatment system of inferior oil, which comprises a conversion reaction unit, an extraction separation unit, a hydrogenation modification unit, a main dilute phase transport bed reactor and an auxiliary 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 main dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet, and the auxiliary dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet;

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 hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the raw material inlet of the main dilute phase conveying bed reactor, and the raw material inlet of the auxiliary dilute phase conveying bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit.

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 in the prior art, 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 realize the maximum production of high-value chemical raw materials and high-quality fuels.

4. The high-quality hydrogenated light cycle oil and the hydrogenated modified oil obtained by the hydrogenation modification unit can be used as catalytic cracking feed to further produce propylene and high-octane gasoline, so that the pressure of surplus diesel markets and insufficient supply of high-octane gasoline is relieved, and the economic benefit of the catalytic unit can be obviously improved by producing propylene.

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 propylene and 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 main 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 regeneration pipe chute

13 regenerative spool valve 14 line 15 line

16 line 17 line 18 line

19 line 20 large oil-gas line 21 line

22 air distributor 23 line 24 cyclone

25 flue gas pipeline 26 pipeline 27 pair dilute phase transport bed reactor

28 line 29 line 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 pipeline

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 method for treating 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 into the bottom of a main dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a first catalyst to be generated and a first reaction product;

introducing the obtained hydrogenated light cycle oil into a secondary dilute phase transport bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a second spent catalyst and a second reaction product;

and regenerating the first catalyst to be regenerated and the second catalyst to be regenerated, and returning the regenerated first catalyst to the main dilute phase conveying bed reactor and the auxiliary dilute phase conveying bed reactor.

According to the invention, inferior oils are well known to those skilled in the art, for exampleThe low-grade oil 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; the residue refers to a bottom distillate obtained by atmospheric and vacuum distillation of crude oil, the atmospheric distillation bottom distillate is generally called atmospheric residue (generally a fraction with a boiling point of more than 350 ℃), the vacuum distillation bottom distillate is generally called vacuum residue (generally a fraction with a boiling point of more than 500 ℃ or 524 ℃), the residue can be at least one selected from topped crude oil, heavy oil obtained from oil sand bitumen and heavy oil with a primary boiling point of more than 350 ℃, and the topped crude oil refers to oil discharged from the bottom of a primary distillation 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 satisfies a condition selected from the group consisting of an API value of less than 27, a distillation range of greater than 350 ℃ (preferably greater than 500 ℃, more preferably greater than 524 ℃), and asphaltene contentGreater 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 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 present invention, the extraction separation is used for separating the easily processable modified oil in the heavy fraction, and the residue is thrown out or returned to carry out the conversion reaction, the extraction separation can be carried out in an extraction solvent under a certain temperature and a certain pressure, preferably the extraction solvent and the heavy fraction are in countercurrent contact for extraction, and the extraction separation can be carried out in any extraction device, such as 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-C ₅At 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 first reaction product and/or second reaction product is separated, and the separated catalytic light cycle oil is returned to the hydro-upgrading unit for 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), as is known in the artThe person can make a reasonable choice for this. 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 may each independently comprise: 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^-1The volume ratio of hydrogen to oil is 300-3000, preferably 800-1500; the hydroupgrading catalyst includes a hydrorefining catalyst and a hydrocracking catalyst, and for example, any catalyst conventionally used for this purpose in the art may be used or may be produced according to any production method conventionally known in the art, and the amount of the hydroupgrading catalyst used in the step may be referred to conventionally known in the art and is 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 to 60 wt% zeolite, 10 to 80 wt% zeolite, based on the dry weight of the hydrocracking catalyst% of alumina, 1-15% by weight of nickel oxide and 5-40% by weight of tungsten oxide, although other compositions of hydrocracking catalysts may also 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 loading volume ratio of the hydrofining catalyst to the hydrocracking catalyst is 1-5:1, and the hydrofining catalyst is loaded at the upstream of the hydrocracking catalyst according to the flow direction of the reaction materials.

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 hydrogenated light cycle oil, which are sent to different reactors for reaction.

According to the present invention, the catalytic reaction unit may select reaction conditions as desired, for example, the conditions of the catalytic cracking reaction include: the reaction temperature is 500-650 ℃, the reaction time is 1-8 seconds, preferably 1-5 seconds, the weight ratio of the catalytic cracking catalyst to the hydro-upgrading oil is 1-50, preferably 3-15, and the weight ratio of the water vapor to the raw oil is 0.05-0.3: 1; the conditions of the catalytic cracking reaction include: the reaction temperature is 560-750 ℃, the preferable temperature is 580-730 ℃, the reaction time is 1-10 seconds, the preferable time is 2-5 seconds, and the weight ratio of the catalytic cracking catalyst to the hydrogenation light cycle oil is 1-100, the preferable time is 5-30; the feed weight ratio of the hydrogenated modified oil to the hydrogenated light cycle oil is 1: (0.1-0.5).

In one embodiment, when the process of the invention is carried out in a catalytic reaction unit, it is generally carried out as follows: the catalyst enters a pre-lifting section of a main dilute phase conveying bed reactor of the catalytic reaction unit and flows upwards under the action of a pre-lifting medium, the preheated hydroupgrading oil and atomized steam are injected into the bottom of the main dilute phase conveying bed reactor together and contact with a regenerated catalyst to perform catalytic cracking reaction and flow upwards at the same time, and the reacted material flow enters a cyclone separator through a reactor outlet; preheated hydrogenation light cycle oil and atomized steam are injected into the bottom of the secondary dilute phase transport bed reactor together, contact with a regenerated catalyst to perform catalytic cracking reaction and flow upwards at the same time, and the reacted material flow enters a cyclone separator through the outlet of the reactor; leading the separated reaction product out of the device, further separating to obtain propylene, high-octane gasoline, light cycle oil and other fractions, and then leading the light cycle oil to a hydro-upgrading unit for upgrading; the separated spent catalyst enters a regenerator for coke burning regeneration, and the regenerated catalyst with recovered activity returns to the reactor for recycling.

According to the invention, the main and secondary dilute-phase transport bed reactors 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 catalytic reaction unit of the present invention, the spent catalyst and the reaction product are generally separated to obtain the spent catalyst and the reaction product, then the obtained reaction product is subjected to a subsequent separation system (for example, a cyclone separator) to separate fractions such as dry gas, liquefied gas, gasoline and diesel oil, then the dry gas and the liquefied gas are further separated by a gas separation device to obtain ethylene, propylene and the like, and the method for separating ethylene, propylene and the like from the reaction product is similar to the conventional technical method in the art, and the present invention is not limited thereto, and is not described in detail herein.

According to the invention, the method of the invention also preferably comprises: regenerating the spent catalyst; and preferably at least a part of the catalyst returned to the main/secondary reactor is regenerated catalyst, for example, all of the regenerated catalyst can be used.

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.

The catalysts used in the catalytic reaction unit according to the present invention are well known to those skilled in the art and conventional options may be employed in the present invention, for example, the catalytic cracking catalyst and the catalytic cracking catalyst may each independently comprise from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay, on a dry basis and based on the weight of the catalyst; the zeolite is taken as an active component, preferably the zeolite is selected from medium-pore zeolite and/or large-pore zeolite, and preferably 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 silica (S) as a binderiO₂) 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 treatment system of inferior oil, which comprises a conversion reaction unit, an extraction separation unit, a hydrogenation modification unit, a main dilute phase transport bed reactor and an auxiliary 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 main dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet, and the auxiliary dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet;

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 hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the raw material inlet of the main dilute phase conveying bed reactor, and the raw material inlet of the auxiliary dilute phase conveying bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit.

According to the invention, the equipment used for catalytic reaction 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 the technical personnel in the field, and the invention is not described again.

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 hydro-modification unit 40 for hydro-modification through a pipeline 39 and the light cycle oil through a pipeline 26, the hydro-modified oil is sent to the bottom of the main dilute phase transport bed reactor 1 through a pipeline 16, the hydrogenated light cycle oil is sent to the bottom of the auxiliary dilute phase transport bed reactor 27 through a pipeline 18, and other products are led out through a pipeline 41.

The pre-lifting medium enters the bottom of the main dilute phase transport bed reactor 1 through a pipeline 14, the regenerated catalyst from a pipeline 12 enters the main dilute phase transport bed reactor 1 after being regulated by a regeneration slide valve 13, and then moves upwards and quickly along the reactor under the lifting action of the pre-lifting medium, the preheated hydroupgrading oil is injected into the main dilute phase transport bed reactor 1 through a pipeline 16 together with the atomized steam from a pipeline 15, and is mixed with the existing material flow of the main dilute phase transport bed reactor 1, and the hydroupgrading oil generates a catalytic cracking reaction on the hot catalyst and moves upwards and quickly; the pre-lifting medium enters the bottom of the auxiliary dilute phase transport bed reactor 27 through a pipeline 29, the regenerated catalyst from a pipeline 28 enters the auxiliary dilute phase transport bed reactor 27 and then moves upwards in an accelerated manner along the reactor under the lifting action of the pre-lifting medium, the preheated hydrogenation light cycle oil is injected into the auxiliary dilute phase transport bed reactor 27 through a pipeline 18 and the atomized steam from a pipeline 17 and is mixed with the existing material flow of the auxiliary dilute phase transport bed reactor 27, and the hydrogenation light cycle oil is subjected to catalytic cracking reaction on the hot catalyst; 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, one part of the degassed regenerated catalyst is circulated to the bottom of a main dilute phase transport bed reactor 1 through a pipeline 12, the circulation quantity of the catalyst can be controlled through a regeneration slide valve 13, the other part of the degassed regenerated catalyst is circulated to the bottom of a secondary dilute phase transport bed reactor 27 through a pipeline 28, 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.

The inferior oils used in the examples and comparative examples were vacuum residue oils, and their properties are shown in table 1.

Example 1

On a medium-sized device, vacuum residue is fed into a conversion reactor of a conversion reaction unit, and the reaction temperature is 430 ℃, the reaction pressure is 17 MPa, and the volume space velocity is 0.5 hour^-1Under the conditions of hydrogen partial pressure of 15.8 MPa and volume ratio of hydrogen to raw material of 2000, under the action of ammonium molybdate as catalyst making conversion reaction, separating reaction product to obtain heavy fraction (distillation range is greater than 350 deg.C); sending the heavy fraction into an extraction separation unit, and carrying out extraction separation by using n-butane as a solvent under the conditions of temperature of 130 ℃, pressure of 4.0 MPa and weight ratio of the extraction solvent to the heavy fraction of 2.5 to obtain modified oil and residue; sending the modified oil and the catalytic light cycle oil to a hydro-modification unit for hydro-modification, wherein the refining temperature and the cracking temperature are both 380 ℃ and the volume space velocity is 0.5 hour^-1And the hydrogen-oil volume ratio is 1000, and the hydrogen partial pressure is 15 MPa, and the hydrogen-oil volume ratio and the hydrogen partial pressure are sequentially contacted with a hydrofining catalyst and a hydrocracking catalyst (the filling volume ratio is 3.5:1) to carry out hydro-upgrading, and the obtained hydro-upgraded oil and the hydrogenated light cycle oil are sent to a catalytic reaction unit. The reaction conditions of the above reaction unit are shown in tables 2 and 3.

The catalytic reaction was carried out on a combined reactor pilot plant with the catalyst being commercially available under the trade designation CGP. The preheated hydrogenation modified oil enters the bottom of a main dilute phase transport bed reactor, catalytic cracking reaction is carried out under the conditions that the reaction temperature is 515 ℃, the reaction time is 1.5 seconds, the weight ratio of the catalyst to the raw oil is 8, and the weight ratio of the steam to the raw oil is 0.10, and a reaction product and a spent catalyst enter a closed cyclone separator from the outlet of the reactor; the preheated hydrogenation light cycle oil enters the bottom of a secondary dilute phase conveying bed reactor, catalytic cracking reaction is carried out under the conditions that the temperature at the outlet of the reactor is 620 ℃, the reaction time is 1.2 seconds, the weight ratio of the catalyst to the raw oil is 15, and the weight ratio of the steam to the raw oil is 0.05, and a reaction product and a spent catalyst enter a closed cyclone separator from the outlet of the reactor; the reaction product and the spent catalyst are quickly separated, the reaction product is cut in a separation system according to the distillation range, so that propylene, high-octane gasoline, light cycle oil and other fractions are obtained, and the light cycle oil returns to the hydro-upgrading unit for upgrading; the spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped catalyst enters a regenerator and is in contact with air for regeneration; the regenerated catalyst enters a degassing tank to remove non-hydrocarbon gas impurities adsorbed and carried by the regenerated catalyst; the degassed regenerated catalyst returns to the reactor for recycling; the catalytic reactor operating conditions and product distribution are listed in table 4.

From the results of tables 2 to 4, it can be seen that the yield of the upgraded oil was 48.4 wt%, the yield of the residue was 51.6 wt%, the residue was thrown out completely, the yield of propylene in the catalytic reaction unit was 9.02%, the yield of gasoline was about 47.62 wt%, and the octane number was as high as 101.

Example 2

On a medium-sized device, vacuum residue is fed into a conversion reactor, and the reaction temperature is 430 ℃, the reaction pressure is 18 MPa, and the volume space velocity is 0.2 hour^-1Carrying out conversion reaction under the conditions of hydrogen partial pressure of 16 MPa and volume ratio of hydrogen to raw materials of 2000 and taking ammonium molybdate as a catalyst, and separating reaction products to obtain heavy fraction (distillation range is more than 350 ℃); sending the heavy fraction into an extraction separation unit, and carrying out extraction separation by using n-butane as a solvent under the conditions of temperature of 130 ℃, pressure of 4.0 MPa and weight ratio of the extraction solvent to the heavy fraction of 2.5 to obtain modified oil and residue; the residue is sent back to the conversion reaction unit, the recycle ratio of the residue is 0.95, the modified oil and the catalytic light cycle oil are sent to a hydro-modification unit for hydro-modification, the refining temperature and the cracking temperature are respectively 382 ℃ and 383 ℃, and the volume space velocity is 0.5 hour^-1Hydrogen-oil volume ratio of 1000 and hydrogen partial pressure of 15 MPaContacting with a hydrofining catalyst and a hydrocracking catalyst (the filling volume ratio is 3.5:1) to carry out hydro-upgrading, and sending the obtained hydro-upgraded oil and the hydrogenated light cycle oil to a catalytic reaction unit. The reaction conditions and the yields of the main products of the above reaction units are shown in tables 2 and 3.

The catalytic reaction unit was run on a medium scale apparatus, with the catalyst being commercially available under the trade designation CGP. The preheated hydrogenation modified oil enters the bottom of a main dilute phase transport bed reactor, catalytic cracking reaction is carried out under the conditions that the reaction temperature is 515 ℃, the reaction time is 1.5 seconds, the weight ratio of the catalyst to the raw oil is 8, and the weight ratio of the steam to the raw oil is 0.10, and a reaction product and a spent catalyst enter a closed cyclone separator from the outlet of the reactor; the preheated hydrogenation light cycle oil enters the bottom of a secondary dilute phase conveying bed reactor, catalytic cracking reaction is carried out under the conditions that the temperature of the outlet of the reactor is 620 ℃, the reaction time is 1.2 seconds, the weight ratio of the catalyst to the raw oil is 15, and the weight ratio of the water vapor to the raw oil is 0.05, reaction products are cut in a separation system according to the distillation range, so that propylene, high-octane gasoline, light cycle oil and other fractions are obtained, and the light cycle oil returns to a hydrogenation modification unit for modification; the spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped catalyst enters a regenerator and is in contact with air for regeneration; the regenerated catalyst enters a degassing tank to remove non-hydrocarbon gas impurities adsorbed and carried by the regenerated catalyst; the degassed regenerated catalyst returns to the reactor for recycling; the catalytic reactor operating conditions and product distribution are listed in table 4.

As can be seen from the results of tables 2-4, the yield of upgraded oil was as high as 60.5 wt%, the yield of bottoms was 39.5 wt%, wherein the yield of rejection bottoms was only 5.2 wt%, the yield of propylene from the catalytic reaction unit was 10.15%, the yield of gasoline was about 53.48 wt%, and the octane number was as high as 103.

Comparative example 1

Basically the same as example 1, except that the vacuum residue is not converted in the reaction unit and directly enters the extraction separation unit to obtain the modified oil and residue.

From the results in tables 2-4, it can be seen that the yield of upgraded oil was only 34.2 wt%, the yield of residue was 65.8 wt%, the residue was thrown out completely, the yield of propylene in the catalytic reaction unit was 6.84%, the yield of gasoline was about 42.76 wt%, and the octane number was 93.4.

Comparative example 2

Essentially the same as example 2 except that the vacuum residue and residuum (from comparative example 1) were not passed through the conversion reaction unit and were directed to the extractive separation unit to yield upgraded oil and residuum.

From the results of tables 2-4, it can be seen that the yield of upgraded oil was only 25.1 wt%, the yield of residue was as high as 74.9 wt%, the yield of propylene from the catalytic reaction unit was 6.32%, the yield of gasoline was about 51.10 wt%, and the octane number was 93.8.

The results of the examples show that the method of the invention greatly improves the yield of the modified oil of the inferior oil, improves the quality of the raw material of the catalytic reactor, and has the obvious advantage of high yield of the high-octane gasoline.

The preferred embodiments of the present invention have been described in detail, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the content of the present invention as long as it does not depart from the gist of the present invention.

TABLE 1

Name (R)	Vacuum residuum
		Density (20 ℃ C.)/(kg/m)3)	1060.3
Degree of API	1.95
		Carbon residue value/weight%	23.2
Element content/weight%
		Carbon (C)	83.87
Hydrogen	9.98
		Sulfur	4.90
Nitrogen is present in	0.34
		Oxygen gas	/
Four components composition/weight%
		Saturation fraction	9.0
Aromatic component	53.6
		Glue	24.4
Asphaltenes	12.7
		Metal content/(microgram/gram)
Ca	2.4
		Fe	23.0
Ni	42.0
		V	96.0
>524 ℃ component content/weight%	100

TABLE 2

TABLE 3

TABLE 4

Claims (11)

1. A method of treating 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 into the bottom of a main dilute phase conveying bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a first catalyst to be generated and a first reaction product;

introducing the obtained hydrogenated light cycle oil into a secondary dilute phase transport bed reactor to contact with a catalytic cracking catalyst and perform catalytic cracking reaction to obtain a second spent catalyst and a second reaction product;

regenerating the first catalyst to be regenerated and the second catalyst to be regenerated, and returning the regenerated first catalyst to the main dilute phase conveying bed reactor and the auxiliary dilute phase conveying bed reactor;

the conditions of the catalytic cracking reaction include: the reaction temperature is 500-650 ℃, the reaction time is 1-8 seconds, and the weight ratio of the catalytic cracking catalyst to the hydro-upgrading oil is 1-50;

the conditions of the catalytic cracking reaction include: the reaction temperature is 560 ℃ and 750 ℃, the reaction time is 1-10 seconds, and the weight ratio of the catalytic cracking catalyst to the hydrogenation light cycle oil is 1-100;

the feed weight ratio of the hydrogenated modified oil to the hydrogenated light cycle oil is 1: (0.1-0.5).

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 the conversion reaction unit comprises a conversion reactor, the 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 conditions for hydro-upgrading comprise: 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.

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 catalytic cracking catalyst and the catalytic cracking catalyst each independently 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 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.

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

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

10. The process of claim 1, wherein the conversion rate of the conversion reaction is 30-70 wt%, the conversion rate of the conversion reaction is (weight of components with distillation range above 524 ℃ in the poor oil-weight of components with distillation range above 524 ℃ in the conversion 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%.

11. A system suitable for the method for treating the inferior oil according to any one of claims 1 to 10, which comprises a conversion reaction unit, an extraction separation unit, a hydro-upgrading unit, a main dilute-phase transport bed reactor and an auxiliary 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 main dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet, and the auxiliary dilute phase transport bed reactor is provided with a catalyst inlet, a raw material inlet and an oil agent outlet;

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 hydrogenation modified oil outlet of the hydrogenation modification unit is communicated with the raw material inlet of the main dilute phase conveying bed reactor, and the raw material inlet of the auxiliary dilute phase conveying bed reactor is communicated with the hydrogenation light cycle oil outlet of the hydrogenation modification unit.

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