CN114426887A

CN114426887A - Method for processing aromatic-rich distillate oil

Info

Publication number: CN114426887A
Application number: CN202011074713.4A
Authority: CN
Inventors: 鞠雪艳; 习远兵; 王锦业; 牛传峰; 王哲; 丁石; 张锐; 葛泮珠
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2020-10-09
Filing date: 2020-10-09
Publication date: 2022-05-03
Anticipated expiration: 2040-10-09
Also published as: CN114426887B

Abstract

The invention relates to a method for processing aromatic-rich distillate oil, wherein the aromatic-rich distillate oil is cut to obtain a first fraction and a second fraction, the first fraction is sent to a diesel oil hydrogenation unit to obtain a third fraction, and the second fraction is sent to a residual oil hydrogenation unit to obtain a fourth fraction; and the obtained third fraction is sent to a first reaction area of the catalytic cracking unit, the obtained fourth fraction is sent to a second reaction area of the catalytic cracking unit, and the reaction effluent of the catalytic cracking unit is separated to obtain dry gas, liquefied gas, catalytic cracking gasoline, catalytic cracking diesel oil and oil slurry. According to the invention, according to the difference of the reaction characteristics of different polycyclic aromatic hydrocarbons in the aromatic-rich distillate oil, the mode of combined treatment of diesel oil hydrogenation, residual oil hydrogenation and catalytic cracking units is adopted, so that the overall utilization efficiency of the three units is effectively improved, and the economical efficiency of the process for processing the aromatic-rich distillate oil is improved.

Description

Method for processing aromatic-rich distillate oil

Technical Field

The invention relates to a method for processing aromatic-rich distillate oil.

Background

With the stricter environmental regulations, the production of clean gasoline and diesel becomes a problem of increasing attention. Sulfur and polycyclic aromatic hydrocarbon in diesel are key factors for generating pollutants, the state continuously sets stricter standards to limit the content of sulfur and polycyclic aromatic hydrocarbon, and the national IV diesel standard is upgraded to the national V diesel standard, and the sulfur mass fraction is required to be reduced from 50 mu g/g to 10 mu g/g; the polycyclic aromatic hydrocarbon mass fraction is further reduced from 11% to 7% from the national V upgrading to the national VI diesel oil standard. At present, about 1/3 of diesel oil from catalytic cracking diesel oil (LCO) in a diesel oil pool in China has the characteristics of high density, high content of impurities and aromatic hydrocarbons, low cetane number, poor stability and the like, can be used as a blending component of clean diesel oil only by a high-severity hydrogenation process, and has high operation cost and poor economic benefit. With the wide application of high-severity catalytic cracking technology such as isoparaffin-rich catalytic cracking technology (MIP), the content of aromatic hydrocarbon, especially polycyclic aromatic hydrocarbon, in the catalytic diesel oil is further increased, and the cetane number is further reduced.

Meanwhile, as the demand of gasoline is increased due to the continuous increase of the reserved quantity of private cars in China, monocyclic aromatic hydrocarbon is gasoline fraction with high octane number, and most of aromatic hydrocarbon in the catalytic cracking diesel oil is aromatic hydrocarbon (polycyclic aromatic hydrocarbon) with more than two rings, the polycyclic aromatic hydrocarbon in the catalytic cracking diesel oil is effectively utilized, and the gasoline component with high aromatic hydrocarbon and high octane number or the aromatic hydrocarbon product is necessary to be produced. In the prior art, LCO is hydrofined, and the aim is to selectively hydrogenate polycyclic aromatic hydrocarbons in the LCO and keep monocyclic aromatic hydrocarbons while saturating the polycyclic aromatic hydrocarbons. In the actual production process, when the catalytic cracking unit and the diesel hydrogenation unit are cooperatively processed, the polycyclic aromatic hydrocarbon content in the hydrocatalytically cracked diesel needs to be limited in order to avoid influencing the catalytic cracking conversion rate. However, in the diesel hydrogenation process, the difficulty of the process of generating monocyclic aromatic hydrocarbon from polycyclic aromatic hydrocarbon, particularly tricyclic aromatic hydrocarbon, is high, and high device operation severity is required, so that the operation period of the diesel hydrogenation device is reduced; moreover, the operating severity increases, which can cause excessive hydrogenation saturation of bicyclic aromatic hydrocarbons to form naphthenic hydrocarbons. There is therefore a need to further optimize the processing scheme for catalytic cracking diesel and other aromatic rich distillates.

CN 108690655A discloses a method for removing polycyclic aromatic hydrocarbons in diesel oil fraction, cutting diesel oil raw oil into light diesel oil fraction and heavy diesel oil fraction, contacting the heavy diesel oil fraction with hydrofining catalyst in a first reaction zone, and carrying out hydrodesulfurization, hydrodenitrogenation and selective hydrodearomatization reaction; the effluent of the first reaction zone enters a second reaction zone to contact a hydrofining catalyst, and further undergoes hydrodesulfurization, hydrodenitrogenation and selective hydrogenation dearomatization under the hydrogenation reaction condition, the effluent of the second reaction zone is separated to obtain a hydrogenated heavy diesel fraction, and the hydrogenated heavy diesel fraction is mixed with a light diesel fraction to obtain a full-cut product, wherein the saturation rate of polycyclic aromatic hydrocarbons in the full-cut product is 85-90%, and the selectivity of monocyclic aromatic hydrocarbons is 80-85%. In the method, the diesel oil hydrogenation unit processes heavy fractions, and the processing difficulty is high, so that the operation period is influenced.

CN 110551525A discloses a method for producing BTX fraction from catalytic cracking diesel oil, which comprises the steps of cutting the catalytic cracking diesel oil into light catalytic cracking diesel oil fraction and heavy catalytic cracking diesel oil fraction, enabling the light catalytic cracking diesel oil fraction to enter a low-pressure hydrocracking unit for reaction, enabling the heavy catalytic cracking diesel oil fraction to enter a hydrotreating unit, enabling obtained liquid phase material flow to enter a catalytic cracking unit and contact with a catalytic cracking catalyst for reaction, enabling a reaction product to enter a fractionation system, and fractionating to obtain dry gas, liquefied gas, BTX-enriched fraction and diesel oil fraction. The method can produce the BTX-rich fraction by adopting a technical scheme of a combined process according to the characteristics of more sulfur and nitrogen impurities and high aromatic hydrocarbon content of the catalytic cracking diesel oil raw material.

CN 108795495 a discloses a method for treating diesel raw material, comprising: cutting the diesel raw oil into light diesel oil fraction and heavy diesel oil fraction, reacting the obtained light diesel oil fraction in a first reaction zone to obtain a component rich in cycloparaffins, reacting the heavy diesel oil fraction in a second reaction zone, allowing the obtained hydrogenated heavy diesel oil fraction to enter a catalytic cracking unit, performing catalytic cracking reaction in the presence of a catalytic cracking catalyst, and separating reaction oil gas to obtain a gasoline product rich in aromatic compounds.

Disclosure of Invention

In the prior art of producing high-octane gasoline components or aromatics by catalytic cracking diesel, in the first step, polycyclic aromatic hydrocarbons in a diesel raw material are selectively hydrogenated to generate monocyclic aromatic hydrocarbons. The inventor of the invention finds that when the content of the polycyclic aromatic hydrocarbon in the hydrocatalytic diesel oil is higher than a certain value, the conversion rate of the hydrocatalytic diesel oil in a catalytic cracking device and the selectivity of the hydrocatalytic diesel oil for producing gasoline can be obviously influenced. Therefore, in an industrial catalytic cracking unit, polycyclic aromatic hydrocarbons of hydrocatalytically cracked diesel fraction are generally limited to ensure higher conversion rate of the unit.

If a higher polycyclic aromatic hydrocarbon saturation rate is to be ensured in the hydrogenation process of the heavy fraction of the diesel oil, not only naphthalene substances which are easy to react in the heavy fraction of the diesel oil need to be hydrogenated and saturated, but also fluorene substances, acenaphthene substances, acenaphthylene substances and aromatic hydrocarbon substances above three rings which are difficult to react in the saturated part need to be hydrogenated and saturated, and higher reaction severity is needed. The result of ensuring the saturation rate of the polycyclic aromatic hydrocarbon is that the monocyclic aromatic hydrocarbon is further saturated into naphthenic hydrocarbon, which causes the increase of hydrogen consumption of a diesel hydrogenation device on one hand and the reduction of the content of the monocyclic aromatic hydrocarbon in a hydrogenation product on the other hand. When the hydrogenated product is processed in a catalytic cracker, the yield of gasoline and the yield of aromatics in the catalytic cracker are adversely affected.

Aiming at the problems in the prior art, the invention provides a method for processing aromatic-enriched distillate oil.

The method provided by the invention comprises the following steps:

(1) cutting the aromatic-rich distillate oil to obtain a first fraction and a second fraction, wherein the content of the aromatic hydrocarbon above three rings in the first fraction is not more than 3 mass percent, and the content of the aromatic hydrocarbon above three rings in the second fraction is more than 20 mass percent; the distillation range of the aromatic-rich distillate oil is 150-400 ℃, and the content of total aromatic hydrocarbon is 50-95 mass%;

(2) sending the first fraction to a diesel hydrogenation unit, contacting with a hydrotreating catalyst for reaction, and separating the obtained reaction effluent to obtain a third fraction;

(3) sending the second fraction to a residual oil hydrogenation unit, carrying out residual oil hydrogenation reaction together with a residual oil raw material, and separating the obtained reaction effluent to obtain a fourth fraction;

(4) and (3) delivering the third fraction obtained in the step (2) to a first reaction area of a catalytic cracking unit to contact with a catalytic cracking catalyst for reaction, delivering the fourth fraction obtained in the step (3) to a second reaction area of the catalytic cracking unit, delivering the reactant flow of the first reaction area to the second reaction area of the catalytic cracking unit, carrying out catalytic cracking reaction together with the fourth fraction, and separating the reaction effluent of the second reaction area to obtain dry gas, liquefied gas, catalytic cracking gasoline, catalytic cracking diesel oil and oil slurry.

In one embodiment of the present invention, the aromatic fraction rich in aromatic hydrocarbons contains aromatic hydrocarbons having a bicyclo ring content or higher in an amount of 40 to 80% by mass.

In one embodiment of the invention, the aromatic-rich distillate oil is selected from one or more of catalytic cracking diesel oil, coking diesel oil, ethylene cracking diesel oil and coal tar.

In one embodiment of the present invention, in step (2) of the present invention, the reaction conditions of the diesel hydrogenation unit comprise: hydrogen partial pressure of 3-12MPa, reaction temperature of 220-400 deg.C, hydrogen-oil volume ratio of 300-1600Nm³/m³The liquid hourly space velocity is 0.3-4.0h^-1(ii) a Preferred reaction conditions for the diesel hydrogenation unit include: hydrogen partial pressure of 6-10MPa, reaction temperature of 260-380 deg.C, hydrogen-oil volume ratio of 500-1200Nm³/m³The liquid hourly space velocity is 0.5-2.0h^-1。

In one embodiment of the invention, the saturation rate of polycyclic aromatic hydrocarbons in the liquid product of the diesel hydrogenation unit is not less than 80%, and the monocyclic aromatic hydrocarbon selectivity is not less than 80%; preferably, the saturation rate of the polycyclic aromatic hydrocarbon is not less than 85 percent, and the selectivity of the monocyclic aromatic hydrocarbon is not less than 90 percent; wherein the content of the first and second substances,

polycyclic aromatic hydrocarbon saturation rate (A)_p1-A_p2)/A_p1*100％

Monocyclic aromatic selectivity ═ a_m2-A_m1)/(A_p1-A_p2)*100％

In the formula: a. the_p1Content of polycyclic aromatic hydrocarbons in the liquid feed,% by mass

A_p2Content of polycyclic aromatic hydrocarbons in the liquid product,% by mass

A_m1Content of monocyclic aromatic hydrocarbons in the liquid feed,% by mass

A_m2Content of monocyclic aromatics in the liquid product, mass%.

The polycyclic aromatic hydrocarbon content refers to the sum of mass fractions of aromatic hydrocarbons with more than two rings in mass spectrum composition data obtained by a mass spectrometry (an analysis method SH/T-0606), wherein the aromatic hydrocarbons with more than two rings comprise the bicyclic aromatic hydrocarbons.

The content of monocyclic aromatic hydrocarbon in the invention refers to the mass fraction of monocyclic aromatic hydrocarbon in mass spectrum composition data obtained by mass spectrometry (analysis method SH/T-0606).

And separating the reaction effluent obtained by the diesel oil hydrogenation unit to obtain a third fraction, wherein the initial boiling point of the obtained third fraction is 150-205 ℃.

The hydrotreating catalyst of the present invention is not particularly limited, and any conventional hydrotreating catalyst or hydrotreating catalyst in the art can be used.

In one of the preferred embodiments of the present invention, the hydrotreating catalyst comprises a carrier and an active metal component supported on the carrier; the carrier is selected from one or more of alumina, silica, titanium oxide, magnesium oxide, zirconia, thorium oxide and beryllium oxide;

the active metal components are at least one metal element selected from VIB group and at least one metal element selected from VIII group, the VIB group metal element is molybdenum and/or tungsten, and the VIII group metal element is cobalt and/or nickel;

the content of at least one metal element from group VIB is 1-30 wt% and the content of at least one metal element from group VIII is 3-35 wt%, calculated as oxides and based on the dry weight of the hydrotreating catalyst.

Preferably, the hydrotreating catalyst has a bulk density of from 0.4 to 1.3g/cm³Average particle diameter of 0.08-1.2mm and specific surface area of 100-300m²/g。

In one preferred embodiment of the present invention, the active metal components are two group VIII metals and one group VIB metal, or the active metal components are one group VIII metal and two group VIB metals, the group VIII metal content is 2.2 to 10 wt% calculated as oxide and based on the catalyst, and the group VIB metal content is 14 to 32 wt% calculated as sulfide and based on the catalyst; the preparation method of the catalyst comprises the following steps: introducing a VIB group metal into a carrier and pre-vulcanizing to obtain a catalyst intermediate; the group VIII metal is then introduced into the catalyst intermediate, dried, optionally calcined, and optionally sulfided to yield the hydroprocessing catalyst.

The preferable hydrotreating catalyst of the invention is prepared by introducing a VIB group active metal component and a VIII group active metal component on an alumina forming carrier in sequence, and performing a pre-sulfurization step between the two active metal components, wherein the active metal contained in the obtained hydrotreating catalyst is molybdenum-tungsten-nickel or molybdenum-cobalt-nickel and the like. Wherein, before, simultaneously or after the active metal is introduced, a step of introducing a phosphorus promoter and/or an organic additive can be further included. The preferred hydroprocessing catalyst of the present invention is thus obtained.

In one embodiment of the present invention, the method for introducing the group VIB metal into the carrier is an impregnation method, which comprises impregnating the carrier with a solution containing the group VIB metal, followed by drying and optionally calcining, wherein the drying conditions include: the temperature is 100-300 ℃, and the time is 1-12 hours; the roasting conditions comprise: the temperature is 300-550 ℃ and the time is 1-10 hours.

In one embodiment of the invention, the presulfiding comprises contacting the group VIB metal-containing support with a sulfiding medium, which is a mixed gas comprising hydrogen sulfide and hydrogen or a solution comprising sulfide and an organic solvent, under conditions comprising: the pressure is 1-15 MPa and the temperature is 300-450 ℃.

Preferably, H in the mixed gas₂The volume fraction of S is 5-20%; the mass fraction of sulfide in the sulfide and organic solvent-containing solution is 1-10 wt%; the organic solvent is cyclohexane and C₆-C₁₀One or more of normal paraffin, kerosene and straight-run diesel oil.

In one embodiment of the invention, the method of introducing the group VIII metal into the catalyst intermediate is an impregnation method comprising impregnating the catalyst intermediate with a group VIII metal-containing solution and then drying, the drying conditions comprising: the temperature is 100-300 deg.C, and the time is 1-12 hr.

The present invention does not particularly require the manner of introduction, provided that it is sufficient to introduce the active metal component to the support and control the amount of introduction, and preferably, the introduction may be carried out by an impregnation method. Specifically, it comprises preparing a solution, preferably an aqueous solution, containing the active metal, and then impregnating the support or catalyst intermediate with the solution, followed by drying, with or without calcination. The active metal can be introduced by one-time impregnation or can be impregnated for multiple times, and when the impregnation is carried out for multiple times, the active metal is dried, roasted or not roasted after each impregnation. After the introduction of the group VIB active metal, calcination is preferred; after introduction of the group VIII active metal, it is preferably not calcined.

In a particular embodiment, the drying and calcining conditions are conventional, for example, drying temperatures of from 100 to 300 ℃, preferably from 100 to 280 ℃, drying times of from 1 to 12 hours, preferably from 2 to 8 hours; the roasting temperature is 300-550 ℃, preferably 300-400 ℃, and the roasting time is 1-10 hours, preferably 2-8 hours.

The aqueous solution containing the active metal component is an aqueous solution containing an active metal compound, and the nickel-containing compound is selected from one or more soluble nickel-containing compounds, for example, one or more of nickel nitrate, nickel acetate, basic nickel carbonate, nickel chloride and soluble nickel complex compounds, and preferably basic nickel carbonate and nickel acetate; the cobalt-containing compound is selected from one or more soluble cobalt-containing compounds, for example, one or more of cobalt nitrate, cobalt acetate, basic cobalt carbonate, cobalt chloride and soluble cobalt complex compounds, preferably basic cobalt carbonate and cobalt acetate; the tungsten-containing compound is one or more selected from soluble compounds of tungsten, such as tungstate and metatungstate, preferably ammonium metatungstate; the molybdenum-containing compound is one or more selected from soluble compounds of molybdenum, such as molybdate and paramolybdate, preferably ammonium paramolybdate.

In one embodiment of the present invention, the hydrotreating catalyst may contain a phosphorus promoter or an organic additive, or may contain both a phosphorus promoter and an organic additive, and preferably contains both a phosphorus promoter and an organic additive. When both the phosphorus adjuvant and the organic additive are contained, the preparation method comprises the step of introducing the phosphorus adjuvant and the organic additive.

When the hydrotreating catalyst contains a phosphorus assistant, the introduction method of phosphorus can be any existing method, for example, a phosphorus-containing compound can be directly introduced when an alumina carrier is prepared and molded; phosphorus and the hydrogenation active metal can be respectively introduced into the carrier, for example, a phosphorus-containing compound solution is firstly contacted with the carrier and is roasted, the roasting temperature is 250-600 ℃, the roasting temperature is 350-500 ℃, the roasting time is 2-8 hours, the roasting time is 3-6 hours, and then the phosphorus and the hydrogenation active metal are contacted with a solution containing the hydrogenation active metal component; furthermore, the phosphorus promoter may be introduced before, simultaneously with, or after the introduction of the group VIII metal component, preferably simultaneously with the introduction of the group VIII metal component. Specifically, an aqueous solution containing a group VIII metal component and a phosphorus-containing compound is prepared, and then the carrier is impregnated. The phosphorus-containing compound is selected from one or more of phosphoric acid, phosphorous acid, phosphate and phosphite, and phosphoric acid is preferred.

In one embodiment of the invention, when the hydrotreating catalyst contains an organic additive. The organic additive is one or more selected from oxygen-containing or nitrogen-containing organic compounds, and the preferable oxygen-containing organic compound is one or more selected from organic alcohol and organic acid; the preferable nitrogen-containing organic compound is one or more selected from organic amines. Examples of the oxygen-containing organic compound include one or more of ethylene glycol, glycerol, polyethylene glycol (molecular weight: 200 to 1500), diethylene glycol, butanediol, acetic acid, maleic acid, oxalic acid, nitrilotriacetic acid, 1, 2-cyclohexanediaminetetraacetic acid, citric acid, tartaric acid, and malic acid, and examples of the nitrogen-containing organic compound include ethylenediamine, EDTA, and ammonium salts thereof. Accordingly, the catalyst preparation process includes a step of introducing an organic additive, and the method of introducing the organic additive may be any method, such as a method of impregnating the support after separately preparing the organic additive into a solution, or a method of impregnating the support with a solution containing the metal component, drying, calcining or not, and then continuing impregnation with a solution containing the organic additive and drying. The drying may be carried out by a conventional method, and is not particularly limited, and for example, the drying temperature is preferably 100 to 300 ℃ and the drying time is preferably 1 to 12 hours, and more preferably 100 to 250 ℃ and the drying time is 2 to 8 hours. The calcination conditions are also conventional, for example the calcination temperature is from 350 to 550 ℃, preferably from 400 to 500 ℃ and the calcination time is from 1 to 10 hours, preferably from 2 to 8 hours. In particular, the organic additive may be introduced before or simultaneously with or after the introduction of the group VIII metal component and/or the group VIB metal component, preferably the additive is introduced simultaneously with the introduction of the group VIII metal component or the group VIB metal component.

In step (3) of the present invention, the present invention has no limitation on the reactor form of the residue hydrogenation unit, the hydrogenation catalyst, or the grading packing manner of the hydrogenation catalyst, and is preferably a fixed bed residue hydrogenation unit.

In one embodiment of the invention, the reaction conditions of the residuum hydrogenation unit include: hydrogen partial pressure of 12-20MPa, reaction temperature of 360-430 deg.C, hydrogen-oil volume ratio of 700-1600Nm³/m³The volume space velocity of the optimal liquid is 0.1-0.6h^-1(ii) a Preferred reaction conditions for the residuum hydrogenation unit include: the hydrogen partial pressure is 13-18MPa, the reaction temperature is 370-420 ℃, and the volume ratio of hydrogen to oil is 800-1300Nm³/m³When it is liquidThe volume space velocity is 0.2-0.5h^-1。

According to the invention, the second fraction rich in aromatic hydrocarbons with double rings and more than three rings is mixed with the residual oil raw material, on one hand, the polycyclic aromatic hydrocarbons are dissolved with the residual oil, so that the viscosity of the residual oil mixed raw material is effectively reduced, the activity and the stability of the residual oil hydrogenation catalyst are fully exerted, and the operation period of the whole residual oil hydrogenation unit is prolonged; and on the other hand, the second fraction is reacted under the condition of a residual oil hydrogenation unit to obtain a fraction rich in monocyclic aromatic hydrocarbon.

In the invention, the fourth fraction is obtained after the reaction effluent of the residual oil hydrogenation unit is separated, and the initial boiling point of the obtained fourth fraction is 220-280 ℃.

In the catalytic cracking unit in the step (4), the third fraction obtained in the step (2) is sent to a first reaction zone of the catalytic cracking unit to contact with a catalytic cracking catalyst for reaction, the fourth fraction obtained in the step (3) is sent to a second reaction zone of the catalytic cracking unit, and a reactant flow in the first reaction zone enters the second reaction zone of the catalytic cracking unit to perform catalytic cracking reaction together with the fourth fraction.

In one embodiment of the invention, the reaction conditions in the first reaction zone of the catalytic cracking unit include: the reaction temperature is 550-600 ℃, and the weight hourly space velocity is 1-600h^-1The reaction pressure is 0.10-1.0MPa, and the weight ratio of the catalytic cracking catalyst to the oil material in the first reaction zone is 6-55: 1; the reaction temperature in the second reaction zone of the catalytic cracking is 15-100 ℃ higher than that in the first reaction zone, preferably 20-60 ℃.

In a preferred case, the first reaction zone of the catalytic cracking unit is a riser reactor, and the second reaction zone of the catalytic cracking unit is a fluidized bed reactor.

In one embodiment of the present invention, the feed to the second reaction zone of the catalytic cracking unit comprises an optional catalytic cracking feedstock, which may be, but is not limited to, one or more of coker gas oil, atmospheric resid, vacuum distillate, vacuum resid.

The catalytic cracking catalyst of the present invention is not limited at all, and the catalytic cracking catalyst may be a catalytic cracking catalyst that is conventional in the art.

According to the present invention, the products obtained by catalytic cracking generally include dry gas, liquefied gas, catalytic cracking gasoline, catalytic cracking diesel oil and slurry oil. The distillation range of the catalytic pyrolysis gasoline is 65-180 ℃, the distillation range of the catalytic pyrolysis diesel oil is 150-450 ℃, preferably 150-400 ℃, the content of total aromatic hydrocarbon is 60-90 wt%, and the content of aromatic hydrocarbon above double rings is 40-80 wt%.

In one embodiment of the invention, the catalytic cracking diesel fraction is directly cut in a fractionation system of the catalytic cracking unit to obtain a first fraction and a second fraction.

In one embodiment of the invention, a fractionation unit is provided, and the catalytically cracked diesel oil and the optional aromatic-rich distillate oil obtained from the catalytic cracking unit are cut to obtain a first fraction and a second fraction.

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

The invention has the following advantages:

1. according to the invention, according to the difference of the reaction characteristics of different types of polycyclic aromatic hydrocarbons contained in the aromatic-rich distillate oil, the diesel oil hydrogenation unit, the residual oil hydrogenation unit and the catalytic cracking unit are effectively combined, and the aromatic-rich distillate oil is cut according to the ring number of the aromatic hydrocarbons and the three units have synergistic effect, so that the overall utilization efficiency of the three units is effectively improved, and the economical efficiency of the process for processing the aromatic-rich distillate oil is improved.

2. The first fraction rich in the monocyclic aromatic hydrocarbon and the bicyclic aromatic hydrocarbon is introduced into the diesel hydrogenation unit, so that the bicyclic aromatic hydrocarbon can generate the monocyclic aromatic hydrocarbon under higher selectivity under mild reaction conditions, the hydrogen consumption is reduced, and the operation period of the diesel hydrogenation unit is greatly prolonged. In particular, the preferred hydrotreating catalyst of the present invention can further reduce hydrogen consumption.

3. The second fraction rich in the aromatic hydrocarbons with double rings and more than three rings is introduced into a residual oil hydrogenation unit, and the fraction rich in the monocyclic aromatic hydrocarbons is obtained under the condition of the residual oil hydrogenation unit. In addition, the viscosity of the residual oil hydrogenation feed is effectively reduced, and the activity and the stability of the residual oil hydrogenation catalyst can be fully exerted, so that the operation period of the whole residual oil hydrogenation unit is prolonged.

4. The catalytic cracking unit adopts zoned feeding, so that the yield of the low-carbon olefin is improved, a high-octane gasoline component can be produced, and the economy of the catalytic cracking unit is improved.

Drawings

FIG. 1 is a schematic flow diagram of a method for processing aromatic-rich distillate oil provided by the invention.

Detailed Description

The method according to the invention will be further illustrated below with the aid of the accompanying drawings, without thereby restricting the invention.

FIG. 1 is a schematic flow chart of one embodiment of the method for processing the aromatic-rich distillate oil provided by the invention. As shown in fig. 1, the aromatic-rich distillate oil 1 is introduced into a fractionation system 2 and separated into a first fraction 3 and a second fraction 23.

The first fraction 3 is introduced into a diesel hydrogenation unit 5 to be in contact reaction with a hydrotreating catalyst 4, wherein bicyclic aromatics are selectively hydrogenated and saturated to generate monocyclic aromatics, a gas-liquid mixture 6 rich in the monocyclic aromatics generated by the reaction passes through a separation device 7 to obtain a third fraction 13, and the third fraction 13 enters a first reaction zone 14 of a catalytic cracking unit to be reacted. Meanwhile, a part of the regenerated catalyst from the catalyst regeneration unit 17 is introduced into the first reaction zone 14 through a line to take part in the reaction.

The second fraction 23 and the residual oil raw material 8 are introduced into a fixed bed residual oil hydrogenation unit 9 to sequentially contact and react with a hydrogenation protective agent, a demetallizing agent and a hydrodesulfurization agent in a residual oil hydrogenation catalyst system 10, and the obtained gas-liquid mixture passes through a separation device 11 to obtain a fourth fraction 12 which enters a second reaction zone 15 of a catalytic cracking unit to react;

in the catalytic cracking unit, the material (including oil gas and catalyst) obtained from the outlet of the first reaction zone 14 enters the second reaction zone 15, and is subjected to catalytic cracking reaction together with the fourth fraction. Introducing materials (including oil gas and catalyst) obtained at the outlet of the second reaction zone 15 into a settler 16 for separation to respectively obtain oil gas and spent catalyst, and introducing the separated spent catalyst into a catalyst regeneration unit 17 for regeneration reaction. The oil gas separated in the settler 16 is introduced into a further catalytic cracking separation unit 18 for separation to obtain dry gas 19, liquefied gas 20, catalytic cracking gasoline 21, catalytic cracking diesel oil 1 and catalytic cracking slurry oil 22.

The process of the present invention will now be further illustrated by the following examples, without thereby limiting the invention.

In the residue hydrogenation unit of the example, the loading hydrogenation catalyst was: the commodity number of the hydrogenation protective agent is RG-20B, the commodity number of the hydrogenation demetallization catalyst is RDM-32, the commodity number of the hydrogenation desulfurization catalyst is RMS-30, and the hydrogenation protective agent and the hydrogenation demetallization catalyst are all commercially available industrial catalysts produced by Chang Ling Branch of China petrochemical catalyst.

In the diesel hydrogenation unit of the examples, the hydrotreating catalyst was an RS-2100 catalyst produced by the chinese petrochemical long ridge catalyst division.

The preparative example is illustrative of the preparation of the hydrotreating catalyst in the diesel hydrogenation unit of the example.

The reagents used in the preparation examples were all chemically pure reagents, unless otherwise specified. Sources of pseudoboehmite or alumina for the preparation of the support include: changling dry glue powder P₀Is a pseudo-boehmite produced by China petrochemical catalyst ChangLing division, the dry basis is 0.73, the specific surface area is 300m²The pore volume was 0.97 mL/g.

Vector S₁Is a common industrial alumina produced by China petrochemical catalyst ChangLing division, and has a specific surface area of 260m²The pore volume was 0.71 mL/g.

Preparation example 1

Weighing S₁200.0 g of carrier, using 168.0 ml containing 70.2 g ammonium paramolybdate and 2.7 g phosphoric acid aqueous solution to soak for 1 hour, 120 degrees C drying for 6 hours, 380 degrees C roasting for 3 hours, get catalyst intermediate Z₁. Taking half weight of Z₁Intermediate to contain H₂H with S volume fraction of 10.0%₂Sulfurizing at 380 deg.C under normal pressure for 4 hr to obtain intermediate Z of catalyst₂. Z was further treated with 55.0mL of an aqueous solution containing 7.0 grams of basic nickel carbonate, 5.0 grams of basic cobalt carbonate, and 8.2 grams of phosphoric acid₂Soaking the intermediate for 1 hr, and drying at 100 deg.C for 8 hr under nitrogen to obtain C₁The catalyst was stored under nitrogen atmosphere for further use. C₁Medium NiO, CoO, MoS₂And P₂O₅The contents by weight of (A) are 2.4%, 1.9%, 21.1% and 4.6%, respectively.

Preparation example 2

Weighing 1000 g of Changling dry rubber powder P₀Extruding into clover-shaped strips with the diameter of the circumscribed circle of 1.6 mm, and drying for 8 hours at 110 ℃. Introducing air, heating to 600 ℃ at the speed of 3 ℃/min, and roasting for 3 hours to obtain the aluminum oxide S₂. Weighing S₂200.0 g of carrier is dipped in 180.0 ml of water solution containing 60.5 g of ammonium paramolybdate for 1 hour, dried for 5 hours at 130 ℃ and roasted for 6 hours at 350 ℃ to obtain catalyst intermediate Z₃. Taking half weight of Z₃Intermediate to contain CS₂Sulfurizing cyclohexane solution with mass fraction of 2.0% at 320 deg.C under 6.0 MPa for 6 hr to obtain catalyst intermediate Z₄. Z was further treated with 64.0mL of an aqueous solution containing 6.9 grams of basic nickel carbonate, 2.8 grams of basic cobalt carbonate, 8.4 grams of phosphoric acid, and 9.2 grams of EDTA₄Soaking the intermediate for 1 hr, and drying at 130 deg.C under nitrogen for 4 hr to obtain C₂The catalyst was stored under nitrogen atmosphere for further use.

C₂Medium NiO, CoO, MoS₂、P₂O₅And the organic additive was 2.3%, 1.1%, 17.9%, 4.0% and 6.0% by weight, respectively. The characterization was carried out according to the procedure of preparation 1, with the parameters within the scope of the invention.

Comparative preparation example 1

This comparative example serves to illustrate a reference diesel hydrogenation catalyst and a process for its preparation. One-half of the weight of the Z1 intermediate from example 2-1 was taken and used in 55.0mL of a mixture containing 7.0 grams of basic nickel carbonate and 5.0 grams of basic nickel carbonate

Cobalt carbonate and 8.2 g phosphoric acid in water solution for 1 hour, then, at 100 ℃ and nitrogen gas under the condition of drying for 8 hours, get D1 catalyst, in the nitrogen atmosphere storage standby. NiO, CoO% and MoO in D1₃And P₂O₅The contents by weight of (a) are 2.4%, 1.9%, 19.4% and 4.7%, respectively.

Catalyst sulfidation

Each catalyst prepared as above employs a temperature programmed sulfiding process to convert an oxidized catalyst to a sulfided catalyst. The vulcanization conditions are as follows: the vulcanization pressure is 6.4MPa, and the vulcanized oil contains CS ₂2% by weight of kerosene, the volume space velocity being 2h^-1And the hydrogen-oil ratio is 300v/v, the constant temperature is kept for 6h at 230 ℃/h, then the temperature is increased to 320 ℃ for vulcanization for 8h, and the temperature increase rate of each stage is 10 ℃/h.

Example 1

By adopting the process flow shown in FIG. 1, the catalytic cracking diesel oil is cut into a first fraction (fraction A) rich in monocyclic aromatics and bicyclic aromatics and a second fraction (fraction B) rich in bicyclic aromatics and aromatic hydrocarbons with more than three rings. Fraction A enters a diesel hydrogenation device for reaction, and is in contact reaction with a hydrotreating catalyst (preparation example 1), and the obtained reaction effluent is separated to obtain a third fraction C (the initial boiling point is 180 ℃). And mixing the fraction B with the residual oil raw material E, feeding the mixture into a residual oil hydrogenation device, carrying out contact reaction with a residual oil hydrogenation catalyst to obtain a reaction effluent, and separating the reaction effluent to obtain a fourth fraction D (the initial boiling point is 250 ℃). Introducing the third fraction C into a first reaction zone of a catalytic cracking unit for reaction; and introducing the fourth fraction D into a second reaction area of the catalytic cracking unit for reaction to obtain dry gas, liquefied gas, catalytic cracking gasoline, catalytic cracking diesel oil and oil slurry, wherein the dry gas is separated to obtain ethylene, and the liquefied gas is separated to obtain propylene.

The residual oil hydrogenation units in examples 1-5 have the same process parameters, the weighted average temperature is 375 ℃, the hydrogen partial pressure is 15MPa, and the liquid hourly volume space velocity is 0.2h^-1The volume ratio of hydrogen to oil was 700. Fruit of Chinese wolfberryThe catalysts of the catalytic cracking units of examples 1-5 were identical in type and all were MMC-2, and were produced by the Qilu division of China petrochemical Co.

The properties of the catalytically cracked diesel feedstock, fraction a and fraction B are shown in table 1.

The properties of resid feed E and fourth fraction D are shown in table 2.

The hydrogenation reaction conditions of the diesel hydrogenation unit, the properties of the third fraction obtained, and the reaction products of the catalytic cracking unit are shown in table 3.

The reaction conditions of the catalytic cracking unit are described in table 4.

TABLE 1

TABLE 2 residual feedstock Properties

Raw oil	Residual oil feedstock E	A fourth fraction D
			Density (20 ℃ C.)/(kg/m)³)	1012.8	941.3
Viscosity (100 ℃ C.)/(mm)²/s)	678.20	24.06
			Residual carbon value/%	19.34	5.25
Sulfur mass fraction/%	4.53	0.17
			Nitrogen mass fraction/%	0.24	0.12
Carbon mass fraction/%	84.82	87.51
			Mass fraction of Ni/(μ g/g)	26.0	6.4
V mass fraction/(μ g/g)	95.0	2.3
			Four component mass fraction/%
Saturation fraction	19.2	41.8
			Aromatic component	54.5	31.5
Glue	19.5	16.1
			Asphaltenes (C)₇Insoluble matter)	6.8	0.8

Examples 2 to 4

The process was carried out as in example 1 except that the hydrotreating catalyst of the diesel hydrogenation unit was varied. The product properties obtained are shown in Table 3.

Example 5

The process was carried out as in example 1, except that the process conditions employed were varied and the product properties obtained are shown in Table 3.

TABLE 3

TABLE 4

Comparative example 1

In the comparative example, the catalytic cracking diesel oil is not fractionated, and after all the catalytic cracking diesel oil enters the diesel oil hydrogenation unit, the obtained liquid material enters the first reaction zone of the catalytic cracking unit for reaction. The product properties obtained are shown in Table 5.

Compared with the comparative example 1, the method of the invention has the advantages that the saturation rate of polycyclic aromatic hydrocarbon and the saturation rate of total aromatic hydrocarbon of the diesel hydrogenation unit are obviously improved, the hydrogen consumption is low, the yield of low-carbon olefin (ethylene and propylene) of catalytic cracking products is obviously improved, the yield of catalytic cracking gasoline is not changed greatly, but the RON value of the gasoline is obviously improved.

Comparative example 2

The treatment was carried out according to the method of comparative example 1, except that this example is the result data after 2000h of operation. The product properties obtained are shown in Table 5.

Comparative example 3

The catalytic cracking diesel was cut to obtain fraction a and fraction B (see table 1). The fraction A does not enter a diesel hydrogenation unit for reaction; and mixing the fraction B with the residual oil raw material E, and then feeding the mixture into a residual oil hydrogenation unit for reaction. Introducing the fraction A into a first reaction zone of catalytic cracking for reaction; and introducing the fourth fraction D into a second reaction zone of the catalytic cracking device for reaction to obtain a catalytic cracking product. The product properties obtained are shown in Table 5.

Comparative example 4

The catalytic cracking diesel was cut to obtain fraction a and fraction B (see table 1). The fraction A enters a diesel oil hydrogenation unit for reaction, and a third fraction is obtained and is introduced into a first reaction zone for catalytic cracking to react; and directly introducing the fraction B into a second reaction zone of the catalytic cracking device without entering a residual oil hydrogenation unit to react with the liquid product of the residual oil hydrogenation unit to obtain a catalytic cracking product. The product properties obtained are shown in Table 5.

TABLE 5

Condition	Comparative example 1	Comparative example 2	Comparative example 3	Comparative example 4
					Diesel hydrogenation unit
Catalyst for diesel hydrogenation reaction zone	C1	C1	-	C1
					Hydrogenation reaction temperature,. degree.C	350	365	-	345
Partial pressure of hydrogen, MPa	8	8	-	8
					Liquid hourly volume space velocity, h^-1	1	1	-	1.2
Hydrogen to oil ratio, Nm³/m³	800	800	-	700
					Liquid product of diesel hydrogenation unit
Total aromatic hydrocarbons, wt.%	69.8	66.4	-	70
					Monocyclic aromatic content,% by weight	58.3	53.1	-	63.8
Polycyclic aromatic hydrocarbons content, wt.%	11.5	13.3	-	6.2
					Polycyclic aromatic hydrocarbon saturation ratio,%	78.9	75.6	-	86.9
Monocyclic aromatic selectivity,%	83.9	74.9	-	92.4
					Hydrogen consumption%	2.33	2.6	-	1.66
Catalytic cracking unit product
					Ethylene	12.2	12.3	12.6	12.3
Propylene (PA)	18.1	19.5	18.4	18.1
					Catalytic crackingGasoline decomposing agent	33.6	35.2	28.8	29.3
RON value of catalytic pyrolysis gasoline	90.8	89.2	91.0	91.2

Example 6

The diesel fuel treatment was carried out as in example 1, except that this example was run for 2000 hours and the test results data for the specific diesel fuel hydrogenation unit conditions are shown in table 6. The product properties obtained are shown in Table 6.

Comparing example 6 with comparative example 2, it can be seen that, in the manner of the present invention, the diesel hydrogenation unit selectively hydrogenates catalytic diesel under significantly mild operating conditions, and the stability of the diesel hydrogenation unit in example 6 is better as can be seen from the stability examination result after 2000 hours of operation.

TABLE 6

Condition	Example 6	Comparative example 2
			Diesel hydrogenation unit
Hydrogenation reaction temperature,. degree.C	348	365
			Partial pressure of hydrogen, MPa	8	8
Liquid hourly space velocity, h^-1	1.2	1.0
			Volume ratio of hydrogen to oil	600	800
Liquid product of a diesel hydrogenation unit%
			Density (20 ℃ C.)/(g/cm)³)	0.9097	0.9184
Total aromatics/%	70.2	66.4
			Monocyclic aromatic content, mass/%)	64.1	53.1
Polycyclic aromatic hydrocarbon content, mass/%)	6.1	13.3
			Polycyclic aromatic hydrocarbon saturation/%)	87.1	75.6
Monocyclic arene selectivity/%)	92.9	74.9
			Hydrogen consumption%	1.8	2.6

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

Claims

1. A method of processing an aromatic-rich distillate oil, the method comprising:

2. The process according to claim 1, wherein the aromatic fraction rich in aromatic hydrocarbons contains aromatic hydrocarbons having a double ring structure or more in an amount of 40 to 80% by mass.

3. The method according to claim 1, wherein the aromatic-rich distillate oil is selected from one or more of catalytic cracking diesel oil, coking diesel oil, ethylene cracking diesel oil and coal tar.

4. The method of claim 1, wherein the reaction conditions of the diesel hydrogenation unit comprise: hydrogen partial pressure of 3-12MPa, reaction temperature of 220-400 deg.C, hydrogen-oil volume ratio of 300-1600Nm³/m³The liquid hourly space velocity is 0.3-4.0h^-1；

Preferred reaction conditions for the diesel hydrogenation unit include: hydrogen partial pressure of 6-10MPa, reaction temperature of 260-380 deg.C, hydrogen-oil volume ratio of 500-1200Nm³/m³The liquid hourly space velocity is 0.5-2.0h^-1。

5. The method of claim 1, wherein the saturation ratio of polycyclic aromatic hydrocarbons in the liquid product of the diesel hydrogenation unit is not less than 80%, and the monocyclic aromatic hydrocarbon selectivity is not less than 80%; preferably, the saturation rate of the polycyclic aromatic hydrocarbon is not less than 85 percent, and the selectivity of the monocyclic aromatic hydrocarbon is not less than 90 percent; wherein the content of the first and second substances,

polycyclic aromatic hydrocarbon saturation rate (A)_p1-A_p2)/A_p1*100％

Monocyclic aromatic selectivity ═ a_m2-A_m1)/(A_p1-A_p2)*100％

A_p2Content of polycyclic aromatic hydrocarbons in the liquid product,% by mass

A_m1Content of monocyclic aromatic hydrocarbons in the liquid feed,% by mass

A_m2Content of monocyclic aromatics in the liquid product, mass%.

6. The process of claim 1, wherein the hydrotreating catalyst comprises a support and an active metal component supported on the support; the carrier is selected from one or more of alumina, silica, titanium oxide, magnesium oxide, zirconia, thorium oxide and beryllium oxide;

7. The process according to claim 6, wherein the active metal components are two group VIII metals and one group VIB metal, or the active metal components are one group VIII metal and two group VIB metals, the group VIII metal content is 2.2-10 wt.% calculated as oxide and based on the catalyst, the group VIB metal content is 14-32 wt.% calculated as sulfide and based on the catalyst; the preparation method of the catalyst comprises the following steps: introducing a VIB group metal into a carrier and pre-vulcanizing to obtain a catalyst intermediate; the group VIII metal is then introduced into the catalyst intermediate, dried, optionally calcined, and optionally sulfided to yield the hydroprocessing catalyst.

8. The process according to claim 7, wherein the method for introducing the group VIB metal into the support is an impregnation method comprising impregnating the support with a group VIB metal-containing solution, followed by drying and optionally calcining, the drying conditions comprising: the temperature is 100-300 ℃, and the time is 1-12 hours; the roasting conditions comprise: the temperature is 300-550 ℃ and the time is 1-10 hours.

9. The process according to claim 7, characterized in that the presulfiding comprises contacting the group VIB metal-containing support with a sulfiding medium, the sulfiding medium being a mixed gas comprising hydrogen sulfide and hydrogen or a solution comprising sulfides and an organic solvent, the conditions of the contacting comprising: the pressure is 1-15 MPa and the temperature is 300-450 ℃.

10. The method of claim 9, wherein H is in the mixed gas₂The volume fraction of S is 5-20%; the mass fraction of sulfide in the sulfide and organic solvent-containing solution is 1-10 wt%; the organic solvent is cyclohexane and C₆-C₁₀One or more of normal paraffin, kerosene and straight-run diesel oil.

11. The process of claim 7 wherein the method of introducing the group VIII metal into the catalyst intermediate is an impregnation method comprising impregnating the catalyst intermediate with a group VIII metal-containing solution followed by drying, the drying conditions comprising: the temperature is 100-300 deg.C, and the time is 1-12 hr.

12. The method of claim 1, wherein the initial boiling point of the third fraction is 150 to 205 ℃.

13. The method of claim 1, wherein the reaction conditions of the residuum hydrogenation unit comprise: hydrogen partial pressure of 12-20MPa, reaction temperature of 360-430 deg.C, hydrogen-oil volume ratio of 700-1600Nm³/m³The volume space velocity of the optimal liquid is 0.1-0.6h^-1；

Preferred reaction conditions for the residuum hydrogenation unit include: the hydrogen partial pressure is 13-18MPa, the reaction temperature is 370-420 ℃, and the volume ratio of hydrogen to oil is 800-1300Nm³/m³The liquid hourly space velocity is 0.2-0.5h^-1。

14. The method of claim 1, wherein the first boiling point of the fourth fraction is 220 to 280 ℃.

15. The method of claim 1, wherein the reaction conditions in the first reaction zone of the catalytic cracking unit comprise: the reaction temperature is 550-600 ℃, and the weight hourly space velocity is 1-600h^-1The reaction pressure is 0.10-1.0MPa, and the weight ratio of the catalytic cracking catalyst to the oil material in the first reaction zone is 6-55: 1; the reaction temperature in the second reaction zone of the catalytic cracking is 15-100 ℃ higher than that in the first reaction zone, preferably 20-60 ℃.

16. The process of claim 15, wherein the first reaction zone of the catalytic cracking unit is a riser reactor and the second reaction zone of the catalytic cracking unit is a fluidized bed reactor.

17. The process of claim 1, wherein the feed to the second reaction zone of the catalytic cracking unit comprises an optional catalytic cracking feedstock.