CN113462431B

CN113462431B - Method for producing diesel oil and jet fuel

Info

Publication number: CN113462431B
Application number: CN202010245420.1A
Authority: CN
Inventors: 张锐; 习远兵; 鞠雪艳; 张毓莹; 丁石; 刘锋; 张乐
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: 2020-03-31
Filing date: 2020-03-31
Publication date: 2023-05-05
Anticipated expiration: 2040-03-31
Also published as: CN113462431A

Abstract

The invention relates to the field of oil refining and chemical industry, and discloses a method for producing diesel oil and jet fuel, which comprises the following steps: (1) Introducing diesel oil raw oil and hydrogen into a first hydrogenation reaction zone filled with a hydrofining catalyst I to perform a first hydrogenation reaction to obtain a first hydrogenation material flow; (2) Introducing the first hydrogenation material flow and normal first-line raw oil into a second hydrogenation reaction zone filled with a hydrofining catalyst II to carry out a second hydrogenation reaction, so that oil flows through the second hydrogenation reaction zone in an upward manner to obtain a second hydrogenation material flow; (3) The second hydrogenated stream is separated to obtain refined diesel and refined jet fuel. The method provided by the invention can be used for treating high-sulfur straight-run diesel oil or inferior secondary processing diesel oil under a relatively mild condition to produce a clean diesel oil product with the sulfur content of less than 10 mug/g and the aromatic hydrocarbon content of more than two rings of less than 7 weight percent.

Description

Method for producing diesel oil and jet fuel

Technical Field

The invention relates to the field of oil refining chemical industry, in particular to a method for producing diesel oil and jet fuel.

Background

Various distillate oils obtained by distillation of crude oil are generally called straight run distillate (distillate oil obtained by direct distillation), wherein straight run kerosene fraction is mainly used for producing jet fuel, and straight run diesel fraction is mainly used for producing diesel. The straight run kerosene is mainly the fraction at 140-280 deg.C, and the diesel oil fraction is mainly the fraction at 200-350 deg.C.

With the increasing environmental awareness and the increasing strictness of environmental regulations, the production and use of clean fuels is becoming a trend. For the cleanliness of diesel, low sulfur and low aromatics content are key to its cleanliness.

The hydrogenation technology is used as an effective desulfurization, denitrification and dearomatization means and plays an increasingly important role in the production of clean fuels. The hydrotreating technology is also various, and single-stage hydrogenation, single-stage serial hydrogenation, two-stage hydrogenation and other technological technologies are presented in sequence. Most of the reactors in the prior art are adiabatic reactors, and because the reaction of reactants is exothermic, a large temperature rise exists in the reactor, so that the reaction is carried out within a reasonable reaction temperature, the inlet temperature of the reactor is usually designed to be lower, and the temperature of a bed layer is gradually increased along with the progress of the reaction until the outlet temperature of the reactor reaches the highest. As the amount of inferior diesel processed by refineries increases, the quality of diesel feedstock is getting worse, which requires higher reaction temperatures for deep hydrodesulfurization.

Typical diesel hydrofining reactors in industry are adiabatic reactors, i.e. the reaction temperature increases in sequence from the top to the bottom of the reactor. For a straight-run diesel hydrogenation device, the temperature rise of the reactor is 30-40 ℃, and the temperature of the outlet of the reactor reaches 380-400 ℃ at the end of the reaction. At such high reaction temperatures, it is difficult to hydrosaturate the polycyclic aromatic hydrocarbons therein. And the straight run kerosene fraction is subjected to over hydrofining at such high reaction temperature, so that sulfides belonging to the natural antiwear agent are deeply removed. When the catalyst is operated to the end of the reaction, the outlet temperature of the reactor can reach 400 ℃, and the difficult-to-remove sulfur-containing compounds and the polycyclic aromatic hydrocarbon saturation stage are mainly carried out on the catalyst, so that the excessive temperature is unfavorable for the reaction.

On an industrial distillate oil hydrofining device, for example, a hydrofining fixed bed process is adopted to simultaneously process normal first-line and diesel oil raw oil, and under the working condition of producing an ultralow-sulfur diesel oil product, higher reaction temperature is required; with mixed feeds, too much of the monocyclic aromatic hydrocarbon in the diesel fraction is easily driven into the kerosene fraction, which results in reduced smoke points in jet fuel products without reducing the liquid yield of the kerosene fraction.

In the prior art, in the process of producing ultralow-sulfur and low-aromatic diesel by adopting the poor-quality diesel raw material, the problems of high reaction severity, poor catalyst stability and quick deactivation exist, and the whole operation period of the device is short.

Therefore, a production method for stably producing ultra-low sulfur diesel oil under a relatively mild condition for a long period is one of the most urgent demands of the oil refining industry.

Disclosure of Invention

The invention aims to overcome the defect that the quality of a product is poor due to the severe reaction conditions at the middle and end stages of the use of a catalyst in the low-sulfur diesel production process in the prior art.

In order to achieve the above object, the present invention provides a method for producing diesel and jet fuel, the method comprising:

(1) Introducing diesel oil raw oil and hydrogen into a first hydrogenation reaction zone filled with a hydrofining catalyst I to perform a first hydrogenation reaction to obtain a first hydrogenation material flow;

(2) Introducing the first hydrogenation material flow and normal first-line raw oil into a second hydrogenation reaction zone filled with a hydrofining catalyst II to carry out a second hydrogenation reaction, so that oil flows through the second hydrogenation reaction zone in an upward manner to obtain a second hydrogenation material flow;

(3) Separating the second hydrogenated stream to obtain refined diesel and refined jet fuel;

the reaction temperature in the second hydrogenation reaction zone is lower than the reaction temperature in the first hydrogenation reaction zone;

the hydrofining catalyst II contains a carrier, and a first element and a second element which are loaded on the carrier, wherein the first element is molybdenum element and/or tungsten element, and the second element is cobalt element and/or nickel element.

The method provided by the invention can be used for treating high-sulfur straight-run diesel oil or low-quality secondary processing diesel oil, and can be used for producing clean diesel oil products with sulfur content less than 10 mug/g and aromatic hydrocarbon content more than double rings less than 7 wt% under a relatively mild condition, and simultaneously producing products with various indexes meeting the quality standard of No. 3 jet fuel.

Compared with the prior art, the whole catalyst system in the preferred condition in the method has better stability, and the running period of the device is obviously improved. Meanwhile, the method provided by the invention can also overcome the defect that the content of the aromatic hydrocarbon above the double rings of the product exceeds the standard caused by poor stability of the catalyst in the method in the prior art.

The method has low equipment requirement and simple process flow, and can obtain high-quality refined diesel oil products and jet fuel products by arranging two hydrogenation reaction areas on a conventional fixed bed reactor.

According to the method of the invention, the reaction effluent of the first hydrogenation reaction zone and the normal first-line raw oil with lighter fraction are mixed, the feeding temperature of the normal first-line raw oil is 50-80 ℃, so that the reaction temperature of the reactant flow entering the second hydrogenation reaction zone can be reduced, after the reaction temperature is regulated to be proper, the reactant flow flows through the catalyst bed layer of the second hydrogenation reaction zone in an upward way, and the ultra-deep hydrodesulfurization reaction and the further hydrogenation saturation of the aromatic hydrocarbon with more than double rings are completed at a higher airspeed. Thus, the method of the present invention also has the advantage of saving energy.

Drawings

Fig. 1 is a process flow diagram of a method of producing diesel and jet fuel in accordance with a preferred embodiment of the present invention.

Description of the reference numerals

1. Diesel oil raw oil 2, hydrogen gas

3. First hydrogenation reaction zone 4, first hydrogenation stream

5. Normal first line raw oil 6, second hydrogenation reaction zone

7. Second hydrogenation material flow 8 and gas-liquid separator

9. Hydrogen-rich gas 10 and recycle hydrogen compression unit

11. Circulating hydrogen 12, cold hydrogen

13. Liquid phase stream 14, fractionation column

15. Refined diesel 16, refined jet fuel

17. Light hydrocarbon product

Detailed Description

The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.

Unless otherwise indicated, the normal pressure according to the present invention means a standard atmospheric pressure.

Unless otherwise specified, the content of an element in terms of oxide represents the most stable pricing oxide of the element.

Unless otherwise indicated, the pressures described herein are gauge pressures.

As previously described, the present invention provides a method of producing diesel and jet fuels, the method comprising:

According to the present invention, the specific loading modes of the first element and the second element are not particularly limited. Particularly preferably, the loading of the first element and the second element is loading the first element and the second element onto the carrier by an impregnation method.

It is to be noted that the process of the present invention introduces all of the material obtained in the first hydrogenation reaction zone (except for the fixedly packed catalyst) into the second hydrogenation reaction zone.

In the invention, the normal first-line raw oil is mainly a fraction at 140-260 ℃, and has higher vaporization rate compared with the diesel oil fraction. Therefore, after the normal first-line raw oil is mixed with the material flow passing through the first hydrogenation reaction zone, the material flow passes through the catalyst bed layer of the second hydrogenation reaction zone in an upward mode, so that the hydrogen sulfide dissolved in the oil product can be effectively brought into the gas phase. So that the concentration of hydrogen sulfide in the liquid phase stream decreases in the second hydrogenation reaction zone from bottom to top. The concentration of hydrogen sulfide in the liquid-phase material flow of the second hydrogenation reaction zone can be greatly reduced, so that the ultra-deep hydrodesulfurization rate can be improved, and the deep aromatic hydrocarbon hydrogenation saturation of the reactant material flow in the second hydrogenation reaction zone is facilitated.

Compared with the traditional hydrodesulfurization and deep dearomatization method, the method of the invention not only reduces the usage amount of the catalyst, effectively reduces the concentration of liquid-phase hydrogen sulfide in the second hydrogenation reaction zone, and reduces the equipment investment.

Preferably, the feeding weight ratio of the diesel oil raw material oil to the normal first-line raw material oil is 1:0.2-0.8; more preferably 1:0.4-0.6.

Preferably, in the hydrofining catalyst II, the content of the first element is 4 to 60 weight percent in terms of oxide and the content of the second element is 1 to 40 weight percent in terms of oxide based on the total weight of the catalyst. More preferably, in the hydrofining catalyst II, the content of the first element in terms of oxide is 20 to 60 wt% and the content of the second element in terms of oxide is 10 to 40 wt% based on the total weight of the catalyst.

The hydrofining catalyst II of the present invention may further contain an auxiliary element such as P, etc., and the content of the auxiliary element is not particularly limited, and may be a content of an auxiliary conventional in the art, for example, the content of the auxiliary element is 1 to 10% by weight in terms of oxide.

Preferably, in the hydrofining catalyst II, the carrier is selected from at least one of alumina, silica, alumina-silica, titania, magnesia, silica-zirconia, silica-thoria, silica-beryllia, silica-titania, titania-zirconia, silica-alumina-thoria, silica-alumina-titania and silica-alumina-magnesia and silica-alumina-zirconia.

In order to obtain a hydrofinishing catalyst suitable for use in the process of the present invention, which is resistant to deep aromatic saturation with hydrogen sulphide at medium and low pressure, according to a preferred embodiment, the hydrofinishing catalyst II is a catalyst prepared by an operation comprising the steps of:

(a) Contacting a first impregnating solution with a carrier to perform first impregnation treatment, and sequentially drying and roasting solid matters subjected to the first impregnation treatment to obtain a first intermediate;

(b) Contacting a second impregnating solution with the first intermediate to perform a second impregnating treatment, and roasting the solid matters subjected to the second impregnating treatment to obtain a catalyst;

the first impregnating solution is an acidic aqueous solution containing a first element, a second element and an organic complexing agent;

the second impregnating solution is an alkaline aqueous solution containing an organic complexing agent.

The catalyst prepared by the method provided by the invention has higher distillate aromatic hydrocarbon hydrogenation saturation activity.

Preferably, in step (a), the conditions of calcination are controlled such that the carbon content in the first intermediate is from 0.03 to 0.5 wt%, more preferably such that the carbon content in the first intermediate is from 0.04 to 0.4 wt%; it is further preferred that the carbon content in the first intermediate is 0.05 to 0.35% by weight.

Preferably, in step (a), the calcination temperature is 350-500 ℃, preferably 360-450 ℃, and the calcination time is 0.5-8h, preferably 1-6h.

Preferably, in step (b), the calcination is carried out at a temperature of 400-550 ℃, preferably 450-550 ℃, for a time of 0.5-8 hours, preferably 2-4 hours.

Preferably, in step (a) and step (b), the organic complexing agent is the same or different, and is each independently selected from at least one of an organic alcohol, an organic acid, an organic amine, and an organic ammonium salt. Illustratively, the organic complexing agent is selected from at least one of ethylene glycol, oxalic acid, citric acid, ethylenediamine tetraacetic acid, and nitrilotriacetic acid.

More preferably, in step (a) and step (b), the organic complexing agent is each independently selected from C _2-7 Organic amine, C _2-7 At least one of the organic ammonium salts of (a).

According to the present invention, the impregnation method may be an isovolumetric impregnation or a supersaturated impregnation, and the temperature of the impregnation is not particularly limited, and may be any of various temperatures which can be reached by the impregnation liquid, and the impregnation time is not particularly limited, so long as a desired amount of a desired component can be supported, for example: the soaking temperature can be 15-60 ℃, and the soaking time can be 0.5-5h.

Preferably, the pH of the first impregnation liquid is 2-6, more preferably the pH of the first impregnation liquid is 3-5.

Preferably, the pH value of the second impregnating solution is 8-11.

In the first impregnating solution and the second impregnating solution, if the acid-base property of the aqueous solution after adding the salt containing the first element and the second element and the organic complexing agent is not satisfactory, the acid-base property of the aqueous solution can be adjusted by a common method, such as adding an acidic substance or an alkaline substance.

Preferably, in the hydrofining catalyst II, the pore volume with the pore diameter of 2nm-40nm accounts for 75-90% of the total pore volume, and the pore volume with the pore diameter of 100nm-300nm accounts for 5-15% of the total pore volume.

Preferably, in the hydrofining catalyst II, the specific surface area is 50-200m ² /g; the pore volume is 0.2-0.4mL/g, and the average pore diameter is 5-40 nm.

According to a preferred embodiment of the present invention, the hydrofining catalyst II is preferably cylindrical, clover or honeycomb in shape.

In the invention, in order to reduce the reaction severity of processing the ultra-low sulfur diesel oil and reduce the influence of reaction interferents on the hydrogenation ultra-deep desulfurization process, two hydrogenation reaction areas connected in series are adopted, and a hydrofining catalyst I is filled in a first hydrogenation reaction area, and preferably has higher reaction performance of ultra-deep hydrogenation desulfurization. The second hydrogenation reaction zone is filled with the hydrofining catalyst II which is preferable in the invention, and the hydrofining catalyst II has higher low-temperature dearomatization reaction performance.

Preferably, the hydrofining catalyst II of the preferred embodiment of the invention has higher hydrogen utilization rate of hydrodesulfurization reaction and is subjected to H in ultra-deep hydrodesulfurization ₂ The inhibition effect of S is smaller, and the diesel oil ultra-deep hydrodesulfurization and dearomatization can be completed under the lower hydrogen-oil volume ratio.

The hydrofinishing catalyst II of the preferred embodiment of the present invention can be prepared using the same method disclosed in CN108568305a, and the preparation example section of the present invention provides a preferred preparation process for the catalyst of the present invention, by way of example. The inventor of the present invention found that the method provided by the present invention can obviously improve the stability of the device and the product quality and prolong the operation period of the device by cooperating with the catalyst disclosed in CN108568305A as the hydrofining catalyst II of the present invention.

Preferably, the hydrofining catalyst I is the same as or different from the hydrofining catalyst II.

More preferably, the hydrofining catalyst I contains a carrier and a hydrogenation metal active component, wherein the hydrogenation metal active component contains at least one metal element selected from the group consisting of VIB and at least one metal element selected from the group consisting of VIII.

Further preferably, in the hydrofining catalyst I, the metal element of group VIB is molybdenum and/or tungsten, and the metal element of group VIII is cobalt and/or nickel.

Preferably, in the hydrofining catalyst I, the content of the VIB group metal element calculated as oxide is 35 to 75 weight percent, preferably 40 to 65 weight percent based on the total weight of the hydrofining catalyst I; the content of the group VIII metal element is 15 to 35% by weight, preferably 20 to 30% by weight.

The type of carrier in the hydrofining catalyst I is not particularly limited, and various carriers conventional in the art can be used. Illustratively, the carrier in the hydrofining catalyst I is alumina and macroporous molecular sieve (i.e. alumina-macroporous molecular sieve), wherein the alumina is further preferably alumina obtained by roasting a hydrated alumina (aluminum hydroxide) colloid compound.

The catalyst of the invention is presulfided with sulfur, hydrogen sulfide or sulfur-containing feedstock, preferably in the presence of hydrogen at a temperature of 170-360 ℃, either ex-situ or in-situ, to convert it to sulfide form.

Particularly preferably, the reaction temperature in the second hydrogenation reaction zone is from 10 to 80 ℃ lower than the reaction temperature in the first hydrogenation reaction zone; further preferably, the reaction temperature in the second hydrogenation reaction zone is 20-80 ℃ lower than the reaction temperature in the first hydrogenation reaction zone.

According to a preferred embodiment, the conditions in the first hydrogenation reaction zone comprise: the reaction temperature is 320-420 ℃, and the volume space velocity is 1.0-3.0h ^-1 The hydrogen partial pressure is 4.0-10.0MPa, and the hydrogen-oil volume ratio is 100-1000:1; the conditions in the second hydrogenation reaction zone include: the reaction temperature is 240-360 ℃ and the volume space velocity is 2.0-10.0h ^-1 Hydrogen partial pressure of2.0-8.0MPa, and the hydrogen-oil volume ratio is 100-1000:1.

According to another more preferred embodiment, the conditions in the first hydrogenation reaction zone comprise: the reaction temperature is 340-420 ℃, and the volume space velocity is 1.0-2.5h ^-1 The hydrogen partial pressure is 6.0-10.0MPa, and the hydrogen-oil volume ratio is 300-1000:1; the conditions in the second hydrogenation reaction zone include: the reaction temperature is 260-320 ℃ and the volume space velocity is 3.0-10.0h ^-1 The hydrogen partial pressure is 2.0-8.0MPa, and the hydrogen-oil volume ratio is 300-1000:1.

Preferably, in step (3), the conditions of the separation are controlled so that the sulfur content in the refined diesel is less than 10 μg/g and the aromatics content above the bicyclo is less than 7 wt%.

Preferably, in step (3), the conditions of the separation are controlled so that the color of the refined diesel is less than No. 0.5.

Preferably, in step (3), the conditions of the separation are controlled such that the refined jet fuel has a Saiki color > +25.

Preferably, in the diesel oil raw material oil, the content of the aromatic hydrocarbon with the bicyclo ring or more is 10-70 wt%; it is further preferable that the content of the aromatic hydrocarbon of the bicyclo or higher is 20 to 65% by weight.

The diesel feedstock of the present invention may be a diesel fraction produced by an atmospheric and vacuum distillation process or by other processes, or a mixture of diesel fractions produced by different processes.

Illustratively, the diesel feedstock of the present invention may be from an atmospheric distillation unit, a catalytic cracking unit, a coker unit.

Preferably, the nitrogen content in the normal first-line raw oil is 10 to 40. Mu.g/g, more preferably 15 to 40. Mu.g/g.

Illustratively, the normal line feedstock of the present invention may be from an atmospheric distillation unit.

The apparatus for carrying out the process is not particularly limited, and for example, one hydrogenation reactor or a plurality of hydrogenation reactors may be disposed in the first hydrogenation reaction zone, or a plurality of catalyst beds may be disposed in one hydrogenation reactor. Preferably, when a plurality of hydrogenation reactors are provided in the first hydrogenation reaction zone, a heat exchanger is provided between each reactor to adjust the inlet temperature of the individual reactor.

A preferred embodiment of the method of the present invention is described in detail below in conjunction with the process flow shown in fig. 1.

The present invention provides a method for producing diesel and jet fuel, the method comprising:

(1) Introducing the diesel oil raw oil 1 and hydrogen 2 into a first hydrogenation reaction zone 3 filled with a hydrofining catalyst I to perform a first hydrogenation reaction to obtain a first hydrogenation material flow;

(2) Introducing the first hydrogenation material flow 4 and normal first-line raw oil 5 into a second hydrogenation reaction zone 6 filled with a hydrofining catalyst II to carry out a second hydrogenation reaction, so that oil flows through the second hydrogenation reaction zone in an upward manner to obtain a second hydrogenation material flow;

(3) Introducing the second hydrogenation material flow 7 into a gas-liquid separator 8 for separation to obtain a hydrogen-rich gas 9 and a liquid-phase material flow 13, compressing and pressurizing the hydrogen-rich gas 9 by a circulating hydrogen compression unit 10, using the hydrogen-rich gas as circulating hydrogen 11, and introducing part of the circulating hydrogen 11 and the fresh-fed hydrogen 2 into the first hydrogenation reaction zone 3; another portion is introduced as cold hydrogen 12 between the beds of the first hydrogenation zone 3; the liquid stream 13 is passed to a fractionation column 14 for fractionation to yield refined diesel 15, refined jet fuel 16 and possibly light hydrocarbon products 17 in minor amounts.

Compared with the prior art, the method has the following specific characteristics:

(1) According to the method, the normal first-line jet fuel raw material is supplemented by arranging the feed inlet between two adjacent reaction areas, so that the color of the product can be improved under high airspeed, and the mercaptan sulfur is removed;

(2) In the method, as the gasification rate of the normal first-line raw material is higher, hydrogen sulfide dissolved in the diesel fraction can be stripped in the second hydrogenation reaction zone, the concentration of the hydrogen sulfide in the liquid phase is reduced, and the hydrodesulfurization reaction rate is further improved;

(3) The method reduces the load of the raw material heating furnace, reduces the load of the circulating hydrogen compressor, and compared with the conventional hydrofining device, omits a jet fuel hydrogenation device, reduces the device investment, and has simple flow and more energy-saving process.

The invention will be described in detail below by way of examples. In the following examples, unless otherwise specified, all the raw materials used are commercially available.

In the following examples, the composition of the catalyst was calculated from the amount of the feed. The specific surface area of the catalyst and the pore distribution, pore diameter and pore volume of 2nm-40nm are measured by adopting a low-temperature nitrogen adsorption method (meeting the GB/T5816-1995 standard), and the pore distribution, pore diameter and pore volume of 100nm-300nm are measured by adopting a mercury intrusion method. The average pore diameter of the catalyst was calculated according to a cylindrical pore model (average pore diameter=total pore volume×4000/specific surface area).

The sulfur content of the diesel oil raw oil is measured by an X-ray fluorescence instrument of the company XOS, and the testing method comprises the following steps: ASTM-7039;

the sulfur content of the diesel oil product is measured by adopting an EA5000 model instrument manufactured by Yes corporation, and the testing method comprises the following steps: SH-0689; the content of aromatic hydrocarbon is analyzed by near infrared spectroscopy.

Hydrofinishing catalyst I used in the examples was commercially available under the trade designation RS-2100.

Hydrofining catalyst II with the trade marks of RS-1000 and RN-410 in the example are produced by China petrochemical catalyst division.

The reaction temperatures of the first hydrogenation reaction zone and the second hydrogenation reaction zone mentioned in the examples are all the weighted average reaction temperatures of the corresponding individual hydrogenation reaction zones. The calculation formula of the weighted average reaction temperature of the single hydrogenation reaction zone is as follows:

weighted average reaction temperature = Σ (reaction area temperature measurement point weight factor x reaction area temperature measurement point display temperature)

Wherein the weight factors are defined as follows:

(1) The weight of the catalyst from the catalyst bed inlet of each hydrogenation reaction zone to the first layer temperature measuring point is represented by the first layer temperature measuring point;

(2) The weight of the catalyst between two adjacent layers of temperature measuring points in each hydrogenation reaction zone is represented by the temperature measuring point of the upper layer, and the other half is represented by the temperature measuring point of the lower layer;

(3) The weight of the catalyst from the lowest temperature measuring point of the catalyst bed layer to the outlet of the catalyst bed layer is represented by the lowest temperature measuring point;

(4) When a plurality of thermocouples exist at each layer of temperature measuring points, the calculated average value of all thermocouple temperature measuring values of the layer is used as the temperature of the temperature measuring points of the layer.

Preparation example 1

Hydrofining catalyst II was prepared and designated S1.

(1) Commercially available white carbon black (specific surface area 220m ² Per g, average pore diameter of 12.7 nm), and basic nickel carbonate powder, and then roasting at 500 ℃ for 3 hours to obtain the inorganic refractory powder containing nickel.

Wherein the amount of basic nickel carbonate used was 16.0 wt.% based on the nickel (calculated as nickel oxide) in the catalyst.

(2) MoO is carried out ₃ Respectively adding basic nickel carbonate and butanol into aqueous solution containing phosphoric acid, heating and stirring until the basic nickel carbonate and butanol are completely dissolved, and then adding tartaric acid until the basic nickel carbonate and butanol are completely dissolved to obtain impregnation solution containing active metal elements.

Wherein the ratio of the mole number of butanol to the weight of the inorganic refractory component is 0.02, and the weight of tartaric acid is 5% of the weight of the inorganic refractory component.

(3) The impregnating solution is mixed with the inorganic refractory component uniformly and then extruded into strips. The catalyst was dried at 150℃for 8 hours to give an oxidation state catalyst having a particle size of 1.6 mm.

Wherein the impregnation solution is mixed with the nickel-containing inorganic refractory powder in a proportion such that the content of molybdenum oxide in the catalyst is 47.0 wt%, the content of nickel oxide is 25.0 wt%, and P is based on the dry weight of the catalyst and calculated as oxide ₂ O ₅ The content was 8.0% by weight, and the content of the inorganic refractory component was 20.0% by weight.

Roasting the catalyst for 2 hours at 400 ℃ to obtain a first intermediate Z1-S1, wherein the carbon content of the Z1-S1 is shown in Table 1; adding 5 g of EDTA into 150 g of deionized water, adding ammonia water to adjust the pH value of the solution to 10.5, stirring to obtain a clear solution, impregnating Z1-S1 with the solution by a saturated impregnation method for 2h, and roasting at 500 ℃ for 3h to obtain the catalyst S1. The content of the hydrogenation metal active component in terms of oxide based on the total amount of S1 is shown in Table 1.

The specific surface area of the hydrorefining catalyst S1 was 155m ² And/g, pore size distribution is between 2nm and 40nm and between 100nm and 300nm, wherein the proportion of pore volume of between 2nm and 40nm to the total pore volume is 89.3% (wherein the proportion of pore volume of between 2nm and 4nm to the total pore volume is 6.7%), the proportion of pore volume of between 100 and 300nm to the total pore volume is 7.4%, the pore volume is 0.31mL/g, and the average pore diameter is 8.0nm.

Preparation example 2

Hydrofining catalyst II was prepared and designated S2.

(1) The commercial zirconium hydroxide powder (specific surface area 180m ² Per g, average pore diameter of 13.3 nm), and basic nickel carbonate powder, and then roasting at 400 ℃ for 3 hours to obtain the inorganic refractory powder containing nickel.

Wherein the amount of basic nickel carbonate used corresponds to 28.0 wt.% of nickel (calculated as nickel oxide) in the catalyst.

(2) Respectively adding ammonium metatungstate, basic nickel carbonate and glycerol into aqueous solution containing phosphoric acid, heating and stirring until the ammonium metatungstate, the basic nickel carbonate and the glycerol are completely dissolved, and then adding caproic acid until the caproic acid is completely dissolved to obtain impregnation solution containing active metals.

Wherein the ratio of the mole number of the glycerol to the weight of the inorganic refractory component is 0.01, and the weight of the caproic acid is 2.5 percent of the weight of the inorganic refractory component.

(3) The impregnating solution is mixed with the inorganic refractory component uniformly and then extruded into strips. Drying at 180 deg.c for 5 hr to obtain oxidized catalyst of 1.6mm size.

Wherein the impregnation solution and the nickel-containing inorganic refractory powder are mixed in such a ratio that the content of tungsten oxide in the catalyst is 45.0% by weight, based on the dry weight of the catalyst and calculated as oxide, the content of nickel oxide is 32.0% by weight, P ₂ O ₅ The content was 3.0% by weight, and the content of the inorganic refractory component was 20.0% by weight.

Roasting the catalyst for 2 hours at 400 ℃ to obtain a first intermediate Z1-S2, wherein the carbon content of the Z1-S2 is shown in Table 1; 10 g of ethylenediamine is placed in 150 g of deionized water, and the mixture is stirred to obtain a clear solution, and an ammonia solution is added to adjust the pH value to 9.5. And (3) impregnating the catalyst S2 with the solution by adopting a saturated impregnation method for 2 hours, and roasting at 450 ℃ for 6 hours to obtain the catalyst S2. The content of the hydrogenation metal active component in terms of oxide based on the total amount of S2 is shown in Table 1.

The specific surface area of the hydrorefining catalyst S2 was 109m ² And/g, pore size distribution is between 2nm and 40nm and between 100nm and 300nm, wherein the proportion of pore volume of between 2nm and 40nm to the total pore volume is 85.6% (wherein the proportion of pore volume of between 2nm and 4nm to the total pore volume is 6.8%), the proportion of pore volume of between 100nm and 300nm to the total pore volume is 12.3%, the pore volume is 0.29mL/g, and the average pore diameter is 10.6nm.

Preparation example 3

Hydrofining catalyst II was prepared and designated S3.

The same process flow as in preparation example 1 was adopted, except that the addition amount of the raw materials was controlled to be different. Catalyst S3 was obtained. The content of the hydrogenation metal active component in terms of oxide based on the total amount of S3 is shown in Table 1.

The specific surface area of the hydrorefining catalyst S3 was 149m ² And/g, pore size distribution is between 2nm and 40nm and between 100nm and 300nm, wherein the proportion of pore volume of between 2nm and 40nm to the total pore volume is 87.6% (wherein the proportion of pore volume of between 2nm and 4nm to the total pore volume is 7.2%), the proportion of pore volume of between 100nm and 300nm to the total pore volume is 10.3%, the pore volume is 0.30mL/g, and the average pore diameter is 7.6nm.

Example 1

The diesel A is used as raw oil, and is a mixed oil obtained by blending 20 wt% of catalytic diesel with middle eastern high-sulfur straight-run diesel fraction.

The diesel A is boosted and then mixed with hydrogen-containing material flow, enters a hydrogenation reactor, firstly passes through a first hydrogenation reaction zone, contacts with a hydrogenation refining catalyst I to carry out hydrogenation refining, then enters a second hydrogenation reaction zone after being mixed with a normal first line A, contacts with a hydrogenation refining catalyst S1 to carry out deep hydrodesulfurization and dearomatization reaction.

The weight ratio of the diesel oil A to the normal line A is 1:0.4.

the partial pressure of hydrogen at the inlet of the reactor of the first hydrogenation reaction zone is 6.4MPa, the reaction temperature of the first hydrogenation reaction zone is 340 ℃, and the liquid hourly space velocity of the first hydrogenation reaction zone is 1.5h ^-1 。

The reaction temperature of the second hydrogenation reaction zone is 260 ℃, the hydrogen partial pressure at the inlet of the reactor of the second reaction zone is 3.2MPa, and the liquid hourly space velocity of the second hydrogenation reaction zone is 6h ^-1 。

The hydrogen-oil volume ratio of the first hydrogenation reaction zone and the second hydrogenation reaction zone is 500.

The total processed oil volume per unit time was 1.56 relative to the total catalyst packing volume.

After the device is continuously operated for 8400 hours, the reaction temperature of the first hydrogenation reactor is 352 ℃, the reaction temperature of the second hydrogenation reactor is 264 ℃, and other process conditions are unchanged.

And (3) carrying out gas-liquid separation on effluent of the second hydrogenation reactor in a low-pressure separator, wherein a gas phase stream obtained by separation in the low-pressure separator is subjected to desulfurization and then is recycled to the reactor inlet, the obtained liquid phase stream enters a fractionating tower for fractionation, light hydrocarbon fraction is produced at the top of the tower, jet fuel components are extracted from the tower, and the bottom of the tower is an ultralow-sulfur low-aromatic diesel fraction.

The main properties of the product are shown in Table 4.

Comparative example 1

The diesel oil A is used as raw oil, the process flow is a conventional single-stage refining process flow, namely, only one hydrogenation reaction zone is provided, the catalyst is RS-2100, the hydrogen partial pressure at the inlet of the reactor is 6.4MPa, the reaction temperature is 338 ℃, the total processed oil volume is 1.56 relative to the total packed volume of the catalyst, and the hydrogen oil volume ratio is 500.

After 8400h of operation, the reaction temperature of the hydrogenation reactor is 355 ℃, and other process conditions are unchanged.

The main properties of the product are shown in Table 4.

Example 2

The diesel oil B is used as raw material oil, and is a mixed oil obtained by blending 50 wt% of catalytic diesel oil with middle eastern high-sulfur straight-run diesel oil fraction.

The diesel oil B is boosted and then mixed with hydrogen-containing material flow, enters a hydrogenation reactor, firstly passes through a first hydrogenation reaction zone, contacts with a hydrogenation refining catalyst I to carry out hydrogenation refining, then enters a second hydrogenation reaction zone after being mixed with a normal line B, contacts with a hydrogenation refining catalyst S2, and carries out deep hydrodesulfurization and dearomatization reaction.

The weight ratio of the diesel oil B to the normal line B is 1:0.5.

the partial pressure of hydrogen at the inlet of the reactor of the first hydrogenation reaction zone is 6.4MPa, the reaction temperature of the first hydrogenation reaction zone is 353 ℃, and the liquid hourly space velocity of the first hydrogenation reaction zone is 1.2h ^-1 。

The reaction temperature of the second hydrogenation reaction zone is 280 ℃, the hydrogen partial pressure at the inlet of the reactor of the second reaction zone is 4.8MPa, and the liquid hourly space velocity of the second hydrogenation reaction zone is 4h ^-1 。

The hydrogen-oil volume ratio of the first hydrogenation reaction zone and the second hydrogenation reaction zone is 800.

The total processed oil volume per unit time was 1.38 relative to the total catalyst packing volume.

After the device is continuously operated for 8400 hours, the reaction temperature of the first hydrogenation reactor is 366 ℃, the reaction temperature of the second hydrogenation reactor is 282 ℃, and other process conditions are unchanged.

The gas-liquid separation and fractionation scheme for the liquid stream were the same as in example 1.

The main properties of the product are shown in Table 5.

Example 3

Diesel C is used as raw oil, which is catalytic diesel.

The diesel oil C is boosted and then mixed with hydrogen-containing material flow, enters a hydrogenation reactor, firstly passes through a first hydrogenation reaction zone, contacts with a hydrogenation refining catalyst I to carry out hydrogenation refining, then enters a second hydrogenation reaction zone after being mixed with a normal first line A, contacts with a hydrogenation refining catalyst S3 to carry out deep hydrodesulfurization and dearomatization reaction.

The weight ratio of the diesel oil C to the normal line A is 1:0.6.

the partial pressure of hydrogen at the inlet of the reactor of the first hydrogenation reaction zone is 6.4MPa, the reaction temperature of the first hydrogenation reaction zone is 360 ℃, and the liquid hourly space velocity of the first hydrogenation reaction zone is 1.0h ^-1 。

The reaction temperature of the second hydrogenation reaction zone is 300 ℃, the hydrogen partial pressure at the inlet of the reactor of the second reaction zone is 6.4MPa, and the liquid hourly space velocity of the second hydrogenation reaction zone is 3h ^-1 。

The hydrogen-oil volume ratio of the first hydrogenation reaction zone and the second hydrogenation reaction zone is 1000.

The total processed oil volume per unit time was 0.95 relative to the total catalyst packing volume.

After the device is continuously operated for 8400 hours, the reaction temperature of the first hydrogenation reactor is 375 ℃, the reaction temperature of the second hydrogenation reactor is 310 ℃, and other process conditions are unchanged.

The main properties of the product are shown in Table 6.

Example 4

The diesel oil A is boosted and then mixed with hydrogen-containing material flow, enters a hydrogenation reactor, firstly passes through a first hydrogenation reaction zone, contacts with a hydrogenation refining catalyst I to carry out hydrogenation refining, then enters a second hydrogenation reaction zone after being mixed with a normal line A, contacts with a hydrogenation refining catalyst RN-410 to carry out deep hydrodesulfurization and dearomatization reaction.

The weight ratio of the diesel oil A to the normal line A is 1:0.3.

The reaction temperature of the second hydrogenation reaction zone is 270 ℃, the hydrogen partial pressure at the inlet of the reactor of the second reaction zone is 3.2MPa, and the liquid hourly space velocity of the second hydrogenation reaction zone is 6h ^-1 。

After the device is continuously operated for 8400 hours, the reaction temperature of the first hydrogenation reactor is 352 ℃, the reaction temperature of the second hydrogenation reactor is 270 ℃, and other process conditions are unchanged.

The main properties of the product are shown in Table 4.

Example 5

The diesel oil B is boosted and then mixed with hydrogen-containing material flow, enters a hydrogenation reactor, firstly passes through a first hydrogenation reaction zone, contacts with a hydrogenation refining catalyst I to carry out hydrogenation refining, then enters a second hydrogenation reaction zone after being mixed with a normal line B, contacts with a hydrogenation refining catalyst RS-1000, and carries out deep hydrodesulfurization and dearomatization reaction.

The weight ratio of the diesel oil B to the normal line B is 1:0.4.

After the device is continuously operated for 8400 hours, the reaction temperature of the first hydrogenation reactor is 366 ℃, the reaction temperature of the second hydrogenation reactor is 290 ℃, and other process conditions are unchanged.

The main properties of the product are shown in Table 5.

Table 1: properties of hydrofining catalysts S1 and S2

Table 2: properties of hydrofining catalysts RN-410 and RS-1000

Catalyst	MoO ₃ Weight percent	NiO, wt%	WO ₃ Weight percent
				RN-410	26.5	4.4	-
RS-1000	2.3	2.6	26

Table 3: nature of raw oil

Table 4: product properties of example 1, example 4, comparative example 1

Project	Example 1	Example 4	Comparative example 1
				Reaction for 800h
Weighted average reaction temperature in the first hydrogenation reaction zone, DEG C	340	340	338
				Weighted average reaction temperature in the second hydrogenation reaction zone, DEG C	260	270
Properties of Diesel product after 800h of reaction
				Sulfur content, μg/g	6	6	7
The content of the arene with more than two rings is weight percent	4.7	5.7	4.6
				Jet fuel product Properties after 800h of reaction
Sulfur content, μg/g	7	8	6
				Mercaptan sulfur content, μg/g	2	3	2
Total nitrogen content/. Mu.g/g	<1	<1	<1
				Saibute colour/number	30	28	30
Reaction 8400h later process conditions
				Weighted average reaction temperature in the first hydrogenation reaction zone, DEG C	352	352	355
Weighted average reaction temperature in the second hydrogenation reaction zone, DEG C	264	270
				Properties of Diesel product after 8400h of reaction
Sulfur content, μg/g	5	6.1	6.8
				The content of the arene with more than two rings is weight percent	4	5.7	6.8
Jet fuel product Properties after reaction 8400h
				Sulfur content, μg/g	6	7	7
Mercaptan sulfur content, μg/g	<2	<2	<2
				Total nitrogen content/. Mu.g/g	<1	<1	<1
Saibute colour/number	30	27	25

Table 5: product Properties of example 2 and example 5

Table 6: product Properties of example 3

Project	Example 3
		Technological conditions after 800 hours of reaction
Weighted average reaction temperature in the first hydrogenation reaction zone, DEG C	360
		Weighted average reaction temperature in the second hydrogenation reaction zone, DEG C	300
Properties of Diesel product after 800h of reaction
		Sulfur content, μg/g	8
The content of the arene with more than two rings is weight percent	6.5
		Jet fuel product Properties after 800h of reaction
Sulfur content, μg/g	10
		Mercaptan sulfur content, μg/g	5
Total nitrogen content/. Mu.g/g	<1
		Saibute colour/number	30
Reaction 8400h later process conditions
		Weighted average reaction temperature in the first hydrogenation reaction zone, DEG C	375
Weighted average reaction temperature in the second hydrogenation reaction zone, DEG C	310
		Properties of Diesel product after 8400h of reaction
Sulfur content, μg/g	9
		The content of the arene with more than two rings is weight percent	6.8
Reaction 8400h jet Fuel product Properties
		Sulfur content, μg/g	9
Mercaptan sulfur content, μg/g	6
		Total nitrogen content/. Mu.g/g	<1
Saibute colour/number	30

From the above results, it is clear that the method of the present invention can produce ultra-low sulfur diesel products and acceptable jet fuel products having very low sulfur content under relatively mild reaction conditions for diesel stock blended with 20 wt% to 50 wt% of the blended diesel, 100 wt% of the catalytic diesel, and two high sulfur or high nitrogen normally first line stock.

The results show that the method provided by the invention has the effects of better device and product quality stability.

And, it can also be seen from comparison of the above results that the reaction temperature can be lower by adopting the method of the present invention when the sulfur content of the product is the same. Even if the pure catalytic cracking diesel oil is processed, the method can also be used for obtaining low-sulfur and low-aromatic refined diesel oil products and qualified jet fuel products.

In addition, the method of the invention has better stability of the catalyst, and after the same running time, the sulfur content of the diesel oil product of the invention is far lower than that of the product in the corresponding comparative example, and the Sagnac color of the jet fuel product is superior to that of the product in the corresponding comparative example.

The method provided by the invention can be used for producing diesel oil with ultralow sulfur and low aromatic hydrocarbon content and qualified jet fuel under a milder reaction condition, and the catalyst has better stability.

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.

Claims

1. A method of producing diesel and jet fuel, the method comprising:

the hydrofining catalyst II contains a carrier, and a first element and a second element which are loaded on the carrier, wherein the first element is molybdenum element and/or tungsten element, and the second element is cobalt element and/or nickel element;

the hydrofining catalyst II is a catalyst prepared by the following steps:

the second impregnating solution is an alkaline aqueous solution containing an organic complexing agent;

the feeding weight ratio of the diesel oil raw oil to the normal first-line raw oil is 1:0.4-0.6.

2. The process according to claim 1, wherein in the hydrofinishing catalyst II, the first element is present in an amount of 4-60 wt% on an oxide basis and the second element is present in an amount of 1-40 wt% on an oxide basis, based on the total weight of the catalyst.

3. The process according to claim 1, wherein in the hydrofinishing catalyst II, the first element is present in an amount of 20-60 wt% on an oxide basis and the second element is present in an amount of 10-40 wt% on an oxide basis, based on the total weight of the catalyst.

4. The method of claim 1, wherein the support is selected from at least one of alumina, silica, alumina-silica, titania, magnesia, silica-zirconia, silica-thoria, silica-beryllia, silica-titania, titania-zirconia, silica-alumina-thoria, silica-alumina-titania and silica-alumina-magnesia and silica-alumina-zirconia.

5. The method according to claim 1, wherein in step (a), conditions of firing are controlled such that a carbon content in the first intermediate is 0.03 to 0.5 wt%.

6. The method according to claim 1, wherein in step (a) and step (b), the kinds of the organic complexing agents are the same or different, and each is independently selected from at least one of an organic alcohol, an organic acid, an organic amine, and an organic ammonium salt.

7. The method of claim 6, wherein in step (a) and step (b), the organic complexing agent is each independently selected from C _2-7 Organic amine, C _2-7 At least one of the organic ammonium salts of (a).

8. The method of any one of claims 1, 6, 7, wherein the pH of the first impregnating solution is from 2 to 6.

9. The method of any one of claims 1, 6, 7, wherein the pH of the second impregnating solution is 8-11.

10. The process according to claim 1, wherein in the hydrofinishing catalyst II, the pore volume of pore diameter of 2nm to 40nm is 75 to 90% of the total pore volume, and the pore volume of pore diameter of 100nm to 300nm is 5 to 15% of the total pore volume.

11. The process according to claim 1, wherein in the hydrofining catalyst II, the specific surface area is 50-200m ² /g; the pore volume is 0.2-0.4mL/g, and the average pore diameter is 5-40 nm.

12. The process according to claim 1, wherein the hydrofinishing catalyst I comprises a support and a hydrogenating metal active component comprising at least one metal element selected from group VIB and at least one metal element selected from group VIII.

13. The process according to claim 12, wherein in the hydrofinishing catalyst I the group VIB metal element is molybdenum and/or tungsten and the group VIII metal element is cobalt and/or nickel.

14. The process of claim 1, wherein the reaction temperature in the second hydrogenation reaction zone is 10-80 ℃ lower than the reaction temperature in the first hydrogenation reaction zone.

15. The process of claim 1, wherein the reaction temperature in the second hydrogenation reaction zone is 20-80 ℃ lower than the reaction temperature in the first hydrogenation reaction zone.

16. The process of claim 1, wherein the conditions in the first hydrogenation reaction zone comprise: the reaction temperature is 320-420 ℃, and the volume space velocity is 1.0-3.0h ^-1 The hydrogen partial pressure is 4.0-10.0MPa, and the hydrogen-oil volume ratio is 100-1000:1.

17. The process of claim 1, wherein the conditions in the second hydrogenation reaction zone comprise: the reaction temperature is 240-360 ℃ and the volume space velocity is 2.0-10.0h ^-1 The hydrogen partial pressure is 2.0-8.0MPa, and the hydrogen-oil volume ratio is 100-1000:1.

18. The process of claim 17, wherein the conditions in the first hydrogenation reaction zone comprise: the reaction temperature is 340-420 ℃, and the volume space velocity is 1.0-2.5h ^-1 The hydrogen partial pressure is 6.0-10.0MPa, and the hydrogen-oil volume ratio is 300-1000:1.

19. The process of claim 17, wherein the conditions in the second hydrogenation reaction zone comprise: the reaction temperature is 260-320 ℃ and the volume space velocity is 3.0-10.0h ^-1 The hydrogen partial pressure is 2.0-8.0MPa, and the hydrogen-oil volume ratio is 300-1000:1.

20. The process of claim 1, wherein in step (3), the conditions of the separation are controlled such that the sulfur content in the refined diesel is less than 10 μg/g and the aromatics content above bicyclo is less than 7 wt%.

21. The method of claim 1, wherein in step (3), the conditions of the separation are controlled such that the refined jet fuel has a seismographic > + No. 25.

22. The method according to claim 1, wherein the content of the bicyclo-or higher aromatic hydrocarbon in the diesel raw oil is 10 to 70 wt%.

23. The method of claim 22, wherein the bi-or higher aromatic hydrocarbon content in the diesel feedstock is 20-65 wt%.

24. The method of claim 1, wherein the nitrogen content of the normal line feedstock is 10-40 μg/g.

25. The method of claim 24, wherein the normal line feedstock has a nitrogen content of 15-40 μg/g.