CN113563925B

CN113563925B - Method for producing jet fuel

Info

Publication number: CN113563925B
Application number: CN202010351316.0A
Authority: CN
Inventors: 张锐; 鞠雪艳; 丁石; 习远兵; 刘锋
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: 2020-04-28
Filing date: 2020-04-28
Publication date: 2023-03-10
Anticipated expiration: 2040-04-28
Also published as: CN113563925A

Abstract

The invention relates to a straight-run jet fuel hydrogenation method, in particular to a method for producing jet fuel. The method for producing jet fuel comprises the following steps: (1) Mixing raw oil with hydrogen-containing gas to obtain a first mixture; (2) Introducing the first mixture into a liquid-phase hydrogenation reactor for carrying out hydrofining reaction to obtain a second mixture; (3) separating the second mixture; wherein, in the step (1), the amount of the hydrogen-containing gas based on the hydrogen gas contained therein is 1 to 10 times the amount of dissolved hydrogen C in the first mixture. The method for producing the jet fuel can calculate the amount of the dissolved hydrogen of the hydrogen in the oil product according to the oil product property of the raw oil, control the consumption of the hydrogen according to the amount of the dissolved hydrogen, inject the hydrogen at one time, reduce the hydrogen consumption, realize the high-efficiency utilization of the hydrogen, and obtain the qualified jet fuel under the condition of reducing the equipment investment and the operation cost of the device.

Description

Method for producing jet fuel

Technical Field

The invention relates to a method for hydrogenating straight-run jet fuel, in particular to a method for producing jet fuel.

Background

Currently, the air transportation industry in China keeps rapid development, the coal consumption is estimated to exceed three million tons by 2020, and the development of the coal technology is emphasized from the consideration of the safety of national energy strategy. At present, jet fuel (also called aviation kerosene, referred to as aviation kerosene for short) mainly comes from the common first-line fraction of an atmospheric distillation device. Straight runThe main quality problems of jet fuel fractions are that mercaptan exceeds the standard, the acid value and the color need to be improved, and the trickle bed jet fuel hydrofining technology is mainly adopted in the current aviation kerosene fraction refining process, and the technology is carried out at the temperature of 230-300 ℃, the pressure of 0.7-1.6 MPa and the airspeed of 2-5 h ^-1 And the hydrogen-oil volume ratio is 30-100 (v/v), mainly reducing the mercaptan in the straight-run aviation kerosene fraction to less than 20 mu g/g, and simultaneously improving the acid value and the color of the oil product.

The reaction conditions of jet fuel hydrorefining are mild, and the requirement of reaction temperature can be met generally at 220-280 ℃. The distillation range of a normal line cut out from the atmospheric tower is generally 140-260 ℃, the distillation temperature is 180-240 ℃, and the temperature requirement of the feed for hydrofining can be met directly or after heat exchange and mixing with hydrogen. At present, a common line of an atmospheric distillation tower is pumped out and then directly enters a downstream jet fuel hydrogenation device or enters an intermediate storage tank as a raw material, the hydrogenation device is generally provided with a new hydrogen compressor and/or a recycle hydrogen compressor and a reactor heating furnace, and the method has the advantages of high device equipment investment, high operation cost, high device energy consumption and high operation cost.

The distillate oil liquid phase hydrogenation technology is a novel hydrogenation technology, hydrogen is dissolved in raw oil in advance, and hydrogen required by hydrogenation reaction is met through liquid phase mass circulation, so that the influence of hydrogen diffusion mass transfer in the conventional trickle bed hydrogenation reaction is overcome, and the hydrogenation reaction is carried out in a dynamic control area. In the distillate oil liquid phase hydrogenation technology, a hydrogen circulation system is not needed, a liquid phase circulating oil system and a hydrogen-oil static mixer can be added, the reaction process is ensured that hydrogen and oil are in a single phase all the time, and the investment of a hydrogen circulation link in the oil refining process accounts for a large proportion of the cost of the whole process. Aiming at the hydrotreatment of the straight-run kerosene fraction, the liquid phase in the reaction system is in a continuous liquid phase, so that the whole catalyst bed layer is soaked by the reactant flow, the wetting effect of the catalyst can be increased, and the effective utilization rate of the catalyst and the higher reaction rate of the hydrodesulfurization can be ensured.

CN102120934a discloses a cyclic liquid phase hydrogenation method, in which a part of the reactant stream is returned as cycle oil to the inlet of the hydrogenation reactor or between catalyst beds, and the mentioned processing raw materials include: naphtha, kerosene, diesel oil, wax oil and residual oil, but a gas-liquid separation device is still required in the method.

US6123835 discloses a two-phase hydrotreating process which achieves sufficient dissolution of hydrogen in feedstock oil by thoroughly mixing the feedstock oil with hydrogen. Then gas-liquid separation is carried out, so that the liquid phase part enters the reactor to carry out hydrogenation reaction. The hydrocarbon oil at the outlet of the reactor is divided into two parts, one part is used as a diluent to be mixed with the raw oil, and the other part enters a subsequent unit. US6428686 complements US6123835 and indicates that the hydrocarbon oil can be withdrawn from the middle of the reactor, mixed well with hydrogen and returned to the reactor for further reaction. The process can process the raw oil with larger chemical hydrogen consumption, broaden the application range of the process, but the existence of a large amount of circulating oil in the method can influence the contact chance of the fresh raw oil and hydrogen with a catalyst, and reduce the mass transfer and reaction speed of the fresh raw oil and hydrogen.

The hydrorefining process involved in the above mentioned liquid phase hydrogenation process has high hydrogen consumption, and a diluent or a circulating oil is required to increase the carrying amount of hydrogen in the raw oil.

Disclosure of Invention

The invention aims to solve the problems of high hydrogen consumption in the hydrofining process and high equipment investment in the jet fuel production process in the prior art

The inventor of the invention finds that the chemical hydrogen consumption of the jet fuel hydrofining reaction is low, the hydrogenation process condition is mild, no additional micromolecular C1-C4 hydrocarbon is generated in the reaction process, and even if the hydrogen contains micromolecular hydrocarbon, the micromolecular hydrocarbon can be dissolved into the jet fuel fraction in the process of researching the jet fuel hydrofining process. Therefore, the inventor considers that the chemical hydrogen consumption and the dissolved hydrogen amount (namely the total consumption of hydrogen) in the raw oil under the process condition can be calculated according to the hydrocarbon property in the raw oil, so that the make-up amount of the hydrogen in the hydrofining reaction process is matched with the consumption of the hydrogen, the gas-liquid separation of the reacted material flow is not needed, the refined jet fuel product can be obtained only by separating hydrogen sulfide and C1-C4 light hydrocarbon from a fractionation system after pressure reduction, and the C1-C4 hydrocarbon dissolved in the raw oil can be separated from the fractionation tower without influencing the property of the refined jet fuel, and based on the thought, the inventor completes the technical scheme of the invention.

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

(1) Mixing raw oil with hydrogen-containing gas to obtain a first mixture;

(2) Introducing the first mixture into a liquid-phase hydrogenation reactor for carrying out hydrofining reaction to obtain a second mixture;

(3) Separating the second mixture;

wherein in the step (1), the amount of the hydrogen-containing gas based on the hydrogen contained therein is 1 to 10 times the amount of dissolved hydrogen C in the first mixture,

the amount of dissolved hydrogen C is calculated according to formula (1):

C＝0.2×P×(19.8×w1+8.1×w2+4.2×w3)EXP((24.0×w1+14.9×w2+28.4×w3)/RT)％

formula (1)

In formula (1): c is the dissolved hydrogen amount of hydrogen in the raw oil, wt%; p is the reaction pressure of the hydrofining reaction, MPa; w1 is the mass fraction of saturated hydrocarbon in the raw oil, wt%; w2 is the mass fraction of monocyclic aromatic hydrocarbon in the raw oil, wt%; w3 is the mass fraction of aromatic hydrocarbon with more than two rings in the raw oil, wt%; r is a gas constant; t is the reaction temperature of the hydrofining reaction, K.

The method for producing jet fuel provided by the invention can calculate the dissolved hydrogen amount of hydrogen in the raw oil according to the oil property of the raw oil, calculate the total consumption of the hydrogen according to the dissolved hydrogen amount, match the make-up amount of the hydrogen in the hydrofining reaction with the consumption of the hydrogen, inject the hydrogen into the device at one time, do not set up a hydrogen circulating system and do not adopt circulating oil as hydrogen carrying oil, can directly reduce the hydrogen consumption, realize the high-efficient utilization of the hydrogen, and obtain the qualified jet fuel under the condition of reducing the equipment investment and the operating cost of the device.

Drawings

Fig. 1 is a schematic flow diagram of a process for producing jet fuel according to a preferred embodiment of the present invention.

Description of the reference numerals

1. Atmospheric tower 2, raw oil

3. Hydrogen mixer 4, first mixture

5. Hydrogenous gas 6, hydrogenation reactor

7. Hydrogen mixing tank 8, second mixture

9. Fractionation system 10, jet fuel

11. Gas at the top of the tower

Detailed Description

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

In the invention, the hydrogen mixing amount (the total consumption of hydrogen in the hydrofining reaction) is slightly more dissolved hydrogen amount on the basis of the chemical hydrogen consumption amount, so that the hydrogen in the reaction system exists in a supersaturated state, and the hydrogen exists in the reaction system as a dispersed phase, and the existence of the hydrogen ensures that the reaction system is in a stable and higher pressure.

As previously mentioned, the present invention provides a method of producing jet fuel, the method comprising:

(1) Mixing raw oil with hydrogen-containing gas to obtain a first mixture;

(3) Separating the second mixture;

the amount of dissolved hydrogen C is calculated according to formula (1):

C＝0.2×P×(19.8×w1+8.1×w2+4.2×w3)EXP((24.0×w1+14.9×w2+28.4×w3)/RT)％

formula (1)

In formula (1): c is the dissolved hydrogen amount (mass fraction) of hydrogen in the oil product, wt%;

p is the reaction pressure of the hydrofining reaction, MPa;

w1 is the mass fraction of saturated hydrocarbon in the raw oil, wt%;

w2 is the mass fraction of monocyclic aromatic hydrocarbon in the raw oil, wt%;

w3 is the mass fraction of dicyclic and above aromatic hydrocarbons in the raw oil, wt%;

r is a gas constant, J/(mol. Multidot.K);

t is the reaction temperature of the hydrofining reaction, K.

In the invention, the content of hydrocarbons (saturated hydrocarbon, monocyclic aromatic hydrocarbon and aromatic hydrocarbon with more than two rings) in raw oil is measured by an SH/T0606-2005 middle distillate hydrocarbon composition measuring method (mass spectrometry).

The R is 8.314J/(mol x K).

According to the invention, firstly, the dissolved hydrogen amount of hydrogen in the raw oil is calculated according to the oil property of the raw oil, the total consumption of the hydrogen is calculated according to the dissolved hydrogen amount, the supplement amount of the hydrogen in the hydrofining reaction process is matched with the consumption of the hydrogen, the required hydrogen-containing gas is supplemented at one time, the hydrofining reaction is carried out under the condition that a catalyst is arranged in a hydrogenation reactor, the utilization rate of the hydrogen is improved, the use amount of the hydrogen-containing gas is reduced, and the reacted material flow is directly subjected to pressure reduction and then enters a fractionation system.

Preferably, in step (1), the amount of the hydrogen-containing gas based on the hydrogen gas contained therein is 1 to 5 times (for example, may be 1 time, 1.2 times, 2 times, 4 times, 5 times or any value therebetween) the amount of dissolved hydrogen C in the first mixture, and more preferably, the amount of the hydrogen-containing gas based on the hydrogen gas contained therein is 1 to 3 times the amount of dissolved hydrogen C in the first mixture. Therefore, the hydrogen excess is less, so that a hydrogen circulation system is not needed, and a hydrogen circulation compression system required in the conventional hydrotreating process is omitted, so that the energy consumption of the device and the equipment investment operation cost are reduced.

In the invention, the liquid phase hydrogenation reactor is preferably an up-flow liquid phase hydrogenation reactor; according to a preferred embodiment, the first mixture obtained by mixing the hydrogen-containing gas with the raw oil enters the catalyst bed from the bottom of the upflow liquid phase hydrogenation reactor and exits from the top of the upflow liquid phase hydrogenation reactor.

In the present invention, the conditions of the hydrotreating reaction are mild, and in the step (2), it is preferable that the conditions of the hydrotreating reaction include: the reaction temperature is 200 to 300 ℃ (for example, may be 200 ℃, 220 ℃, 240 ℃, 260 ℃, 280 ℃, 300 ℃ or any value between the above values), more preferably 220 to 280 ℃, and still more preferably 260 to 280 ℃; the reaction pressure is 1.0 to 4.0MPa (for example, 1MPa, 2MPa, 3MPa, 4MPa or any value therebetween), and more preferably 2.0 to 4.0MPa; the liquid hourly space velocity is 2 to 10h ^-1 (for example, it may be 2 h) ^-1 、4h ^-1 、6h ^-1 、8h ^-1 、10h ^-1 Or any value therebetween), preferably 3 to 6h ^-1 More preferably 4 to 6 hours ^-1 。

In the present invention, in the step (2), it is preferable that the hydrotreating reaction is carried out in the presence of a hydrotreating catalyst. In order to provide a hydrorefining catalyst with high reactivity and good hydrodesulfurization activity under mild conditions, it is preferable that the hydrorefining catalyst comprises a carrier and a hydrogenation-active metal component supported on the carrier, and the content of the hydrogenation-active metal component is 5 to 50 wt% (for example, 5 wt%, 10 wt%, 15 wt%, 20 wt%, 30 wt%, 40 wt%, 50 wt%, or any value therebetween) in terms of oxide based on the total amount of the hydrorefining catalyst.

In the invention, the hydrogenation metal active component preferably contains at least one of VIB group metal elements and at least one of VIII group metal elements; more preferably, the group VIB metal element is molybdenum and/or tungsten, and the group VIII metal element is cobalt and/or nickel.

According to a particularly preferred embodiment, the content of group VIB metal elements in the present invention, expressed as oxides, is from 4 to 40% by weight (for example, 4%, 8%, 15%, 20%, 30%, 35%, 40% by weight or any value between the aforementioned values), preferably from 8 to 35% by weight, based on the total amount of the hydrofinishing catalyst; the content of the group VIII metal element is 1 to 10% by weight (for example, may be 1% by weight, 2% by weight, 3% by weight, 5% by weight, 10% by weight, or any value therebetween), and preferably 2 to 5% by weight, in terms of oxide.

Illustratively, the hydrofining catalyst of the invention is an RSS-2 catalyst developed by the research institute of petrochemical engineering science, and the catalyst is a NiMo system kerosene hydrofining catalyst.

In the present invention, the hydrogenation catalyst in the oxidation state is preferably sulfided prior to use, the sulfiding scheme being specified or designated by the catalyst supplier, and the invention is not described in detail herein.

In the present invention, kerosene fractions produced by an atmospheric and vacuum distillation process or by other processes, or a mixture of kerosene fractions produced by different processes may be used as the feedstock oil. Preferably, the raw oil is a normal first-line distillate oil obtained by atmospheric distillation of crude oil; further preferably, the sulfur content of the normal first-line distillate oil is 200-5000 mug/g, more preferably 200-3500 mug/g; preferably, the content of aromatics with more than two rings in the normal first-line distillate is not more than 5 wt%, preferably not more than 3 wt%.

In the present invention, in the step (1), the temperature of the feedstock oil before mixing with the hydrogen-containing gas is preferably 200 to 300 ℃. When the raw oil is the first-line distillate oil, the temperature of the first-line distillate oil extracted from the atmospheric tower is 200-300 ℃, and when the temperature is not in the range, the first-line distillate oil can exchange heat with other media to 200-300 ℃, and then enters the hydrogenation reactor for reaction. The common first-line distillate oil is adopted as the raw oil, the combination of the atmospheric tower and the jet fuel hydrogenation process is realized, the heat of the atmospheric tower is effectively utilized, and the equipment investment and the energy consumption of the device are reduced.

In the present invention, the pressure of introducing the hydrogen-containing gas is the same as the pressure of the hydrorefining reaction, so that the hydrorefining reaction can be directly performed after the first mixed gas enters the liquid phase reactor, and preferably, in the step (1), the pressure of introducing the hydrogen-containing gas is 1.0 to 4.0MPa.

Preferably, in the present invention, the hydrogen-containing gas contains 80 to 99% by volume of hydrogen and 1 to 20% by volume of methane, ethane and heavier hydrocarbon gases. More preferably, the hydrogen-containing gas has a volume fraction of hydrogen of 90% to 99%, and methane, ethane and heavier hydrocarbon gases have a volume fraction of 1% to 10%. Further preferably, the hydrogen-containing gas is from at least one of a hydrogen pipeline network of a hydrorefining unit, a hydrocracking unit, a catalytic reforming unit and an oil refinery.

Fig. 1 is a schematic flow chart of a process for producing jet fuel according to a preferred embodiment of the present invention, and the method provided by the present invention is further described below with reference to the accompanying drawings, in which many devices such as pumps, furnaces, heat exchangers, compressors, etc. are omitted, but are well known to those skilled in the art.

As shown in fig. 1, the jet fuel is produced by using a liquid phase hydrogenation system in the invention, wherein the liquid phase hydrogenation system comprises a hydrogen dissolving unit, a hydrogenation unit and a fractionation unit, and in a preferred embodiment of the invention, the jet fuel production process comprises the following specific steps: raw oil 2 (or subjected to heat exchange) obtained after normal-line distillate oil is extracted from a normal pressure tower 1 is mixed with hydrogen-containing gas 5 in a hydrogen mixer 3, hydrogen is dissolved in the raw oil 2 to a saturated state, a first mixture 4 (saturated hydrogen material mixture) is obtained, the first mixture 4 is conveyed to the bottom of a hydrogenation reactor 6 and passes through a hydrofining catalyst bed from bottom to top, the first mixture 4 is subjected to liquid-phase hydrofining reaction under the action of a hydrofining catalyst, mercaptan sulfur is removed, an acid value is obtained, a product color is improved, a second mixture 8 is obtained, the second mixture 8 flows out of the top of the hydrogenation reactor 6 and enters a fractionating system 9 for rectification, overhead gas 11 (hydrogen sulfide and C1-C4 hydrocarbons) is obtained at the top of the tower, and jet fuel 10 is obtained at the bottom of the tower.

In the present invention, it is preferable to design the hydrogen mixing tank 7 between the hydrorefining catalyst beds, depending on the properties of the feedstock, in order to improve the hydrogen distribution in the feedstock and to increase the concentration of hydrogen in the feedstock.

The invention adopts the continuous liquid phase hydrogenation process, so that the whole catalyst bed layer is always soaked by the reactant flow and is in a continuous liquid phase system, the wetting effect of the catalyst is increased, the reaction volume of the catalyst is effectively utilized, the utilization rate of the catalyst is improved, and the efficiency of the hydrofining reaction is improved.

In the present invention, in order to ensure that the hydrogen gas and the feedstock oil are always in a single phase during the reaction, it is preferable that the hydrogen mixer 3 be a commercial static mixer or a disperser of ceramic nanotubes.

In the invention, the temperature of the second mixture 8 is higher when the second mixture flows out of the hydrogenation reactor 6, the second mixture can exchange heat with raw oil or other media after flowing out, then the second mixture is directly reduced to the operating pressure of a fractionation system, and then the second mixture enters the fractionation system for fractionation. The fractionation system has a fractionation column or stripper column with a reboiler to obtain jet fuel.

Preferably, the fractionating tower (or the stripping tower) has the top operation pressure of 250-350 KPa, the bottom operation temperature of 210-220 ℃ and the top operation temperature of 70-110 ℃.

The present invention will be described in detail below by way of examples.

In the following examples, atmospheric tower normal first line distillate oil was used as feed oil, and the properties are shown in Table 1; the hydrofining catalyst is an RSS-2 catalyst developed by the research institute of petrochemical engineering science, and the catalyst is a NiMo system kerosene hydrofining catalyst.

Table 1 main properties of the first-line distillate

Item	Raw oil A	Raw oil B
			Density (20 ℃ C.)/g.cm ^-3	0.8074	0.7967
Total sulfur content/μ g/g	2000	3200
			Mercaptan sulfur content/. Mu.g/g	108	120
The mass fraction w 1/weight% of saturated hydrocarbons	83	79
			Mass fraction w 2/weight% of monocyclic aromatic hydrocarbons	14	18
Quality of aromatic hydrocarbons having more than two ringsFraction w 3/weight%	2.5	3
			Distillation Range (ASTM D-86)/. Deg.C	161～251	170～260

The hydrogen mixer adopts a ceramic membrane pipe material with the average pore diameter of about 50 nanometers produced by Zhongpetrochemical ChangLing division company, and gas-phase hydrogen passes through the ceramic pipe from outside to inside in the radial direction in the mixing device, is uniformly distributed and is quickly dissolved in liquid-phase oil products.

The hydrogen-containing gas comes from a hydrogen catalytic reforming device, and the volume fraction of hydrogen in the hydrogen-containing gas is 92%.

The amount of hydrogen dissolved in the oil, C, is calculated according to formula (1):

C＝0.2×P×(19.8×w1+8.1×w2+4.2×w3)EXP((24.0×w1+14.9×w2+28.4×w3)/RT)％

formula (1)

In formula (1): c is the dissolved hydrogen amount of the hydrogen in the oil product; p is the reaction pressure of the hydrofining reaction; w1 is the mass fraction of saturated hydrocarbon in the raw oil; w2 is the mass fraction of monocyclic aromatic hydrocarbon in the raw oil; w3 is the mass fraction of aromatic hydrocarbon above double rings in the raw oil; r is a gas constant; t is the reaction temperature of the hydrofining reaction.

The second mixture from the hydrogenation reactor was subjected to heat exchange and pressure reduction and then fractionated in a fractionation column, in the following example, a conventional fractionation unit having a reboiler was used, and the main operating conditions of the fractionation unit are shown in table 2.

TABLE 2 main operating conditions of the fractionation system

Item	Examples 1 to 5 parameters	Comparative examples 1-2 parameters
			Column feed temperature/. Degree.C	200	200
Overhead temperature/. Degree.C	120	100
			Temperature of the bottom of the column/. Degree.C	240	240
Overhead reflux drum pressure/KPa	200	180
			Reflux ratio	0.1	0.2

Example 1

In this example 1, jet fuel was produced by referring to the flow shown in FIG. 1, using the feed oil shown in Table 1, the catalyst RSS-2, and the operating parameters of the fractionation column are shown in Table 2. The amount of hydrogen injected was 2 times the amount of dissolved hydrogen, and the other hydrogenation conditions and the properties of the obtained product are shown in Table 3.

Example 2

In this example 1, jet fuel was produced by referring to the flow shown in FIG. 1, using feed oil B shown in Table 1, with RSS-2 as the catalyst, and with the operating parameters of the fractionation column shown in Table 2. The amount of hydrogen injected was 1 time the amount of dissolved hydrogen, and the other hydrogenation conditions and the properties of the obtained product are shown in Table 3.

Example 3

In this example 1, jet fuel was produced by referring to the flow shown in FIG. 1, using feed oil A shown in Table 1, RSS-2 as a catalyst, and the operating parameters of the fractionator shown in Table 2. The amount of hydrogen injected was 1.2 times the amount of dissolved hydrogen, and the other hydrogenation conditions and the properties of the obtained product are shown in Table 3.

Example 4

In this example 1, jet fuel was produced by referring to the flow shown in FIG. 1, using feed oil A shown in Table 1, with RSS-2 as the catalyst, and with the operating parameters of the fractionation column shown in Table 2. The amount of hydrogen injected was 2 times the amount of dissolved hydrogen, and the other hydrogenation conditions and the properties of the obtained product are shown in Table 3.

Example 5

In this example 1, jet fuel was produced by referring to the flow shown in FIG. 1, using feed oil B shown in Table 1, with RSS-2 as the catalyst, and with the operating parameters of the fractionation column shown in Table 2. The amount of hydrogen injected was 4 times the amount of dissolved hydrogen, and the other hydrogenation conditions and the properties of the obtained product are shown in Table 3.

Comparative example 1

A conventional down-flow fixed bed hydrogenation process flow is adopted, a gas-liquid separation tank is arranged behind a reactor, a separated liquid phase is subjected to heat exchange and then enters a fractionation device, raw oil shown in table 1 is adopted, a catalyst is RSS-2, and the operation parameters of a fractionation tower are shown in table 2. The volume ratio of the injected hydrogen to the raw oil (hydrogen-oil ratio) was 100, and the injected hydrogen was calculated to be 16 times the amount of dissolved hydrogen by the method, and the other hydrogenation conditions and the properties of the obtained product are shown in table 4.

Comparative example 2

A conventional downflow fixed bed hydrogenation process flow is adopted, a gas-liquid separation tank is arranged behind a reactor, a separated liquid phase is subjected to heat exchange and then enters a fractionation device, raw oil B shown in table 1 is adopted, a catalyst is RSS-2, and the operation parameters of a fractionation tower are shown in table 2. The volume ratio of the injected hydrogen to the feedstock oil (hydrogen-oil ratio) was 150, and the hydrogen injection amount was calculated to be 12 times the dissolved hydrogen amount by the method, and the other hydrogenation conditions and the properties of the obtained product are shown in table 4.

Table 3 example process conditions and test results

Process conditions	Example 1	Example 2	Example 3	Example 4	Example 5
						Reaction pressure/Mpa	3	4	4	2	2
Reaction temperature/. Degree.C	260	280	280	260	240
						Volume space velocity/h ^-1	4	4	4	4	6
Amount of dissolved hydrogen C	0.063	0.086	0.084	0.042	0.040
						Hydrogen gas injection amount/dissolved hydrogen amount	2	1	1.2	2	4
Raw oil	Raw oil A	Raw oil B	Raw oil A	Raw oil A	Raw oil B
						Properties of refined oils
Density (20 ℃ C.)/g.cm ^-3	0.805	0.801	0.803	0.803	0.801
						Total sulfur content/μ g/g	720	740	500	1000	1800
Mercaptan sulfur content/. Mu.g/g	5	6	3	8	10
						Polycyclic aromatic hydrocarbons content, wt.%	2.4	2.8	2.4	2.5	2.5
Distillation Range (ASTM D-86)/. Deg.C	160～250	170～260	160～250	160～250	160～250

Table 4 example process conditions and test results

Process conditions	Comparative example 1	Comparative example 2
			Reaction pressure, mpa	2	4
Reaction temperature of	260	280
			Volume space velocity, h ^-1	4	4
Volume ratio of hydrogen to oil	100	150
			Amount of dissolved hydrogen C	0.042	0.084
Hydrogen gas injection amount/dissolved hydrogen amount	16	12
			Raw oil	Raw oil A	Raw materialsOil B
Properties of refined oils
			Density (20 ℃ C.)/g.cm ^-3	0.805	0.803
Total sulfur content/μ g/g	800	750
			Mercaptan sulfur content/. Mu.g/g	6	5
Polycyclic aromatic hydrocarbons content, wt.%	2.4	2.7
			Distillation range (ASTM D-86), deg.C	160～250	170～260

The test results of examples 1-5 and comparative examples 1-2 show that the method of the invention does not need to adopt circulating oil as hydrogen carrying oil, only needs hydrogen with slightly more than dissolved hydrogen to be injected into the processed raw oil through a hydrogen mixer to carry out continuous liquid phase hydrofining, can lead the mercaptan sulfur content of the product to be less than 10 mug/g, the polycyclic aromatic hydrocarbon content to be less than 3 weight percent, and other indexes to meet the jet fuel standard of GB 6537-2006, thereby reducing the hydrogen consumption, realizing the high-efficiency utilization of the hydrogen, and obtaining the qualified jet fuel under the condition of reducing the equipment investment and the operation cost of the device.

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 producing jet fuel, comprising:

(1) Mixing raw oil with hydrogen-containing gas to obtain a first mixture;

(2) Introducing the first mixture into a liquid-phase hydrogenation reactor for carrying out hydrofining reaction to obtain a second mixture; the liquid phase hydrogenation reactor is an up-flow liquid phase hydrogenation reactor;

(3) Separating the second mixture; obtaining jet fuel;

wherein in the step (1), the amount of the hydrogen-containing gas based on the hydrogen contained therein is 1 to 10 times of the amount C of the dissolved hydrogen in the first mixture,

the amount of dissolved hydrogen C is calculated according to formula (1):

c =0.2 × P × (19.8 × w1+8.1 × w2+4.2 × w 3) EXP ((24.0 × w1+14.9 × w2+28.4 × w 3)/RT)% formula (1)

In formula (1): c is the dissolved hydrogen amount of hydrogen in the raw oil, wt%; p is the reaction pressure of the hydrofining reaction, MPa; w1 is the mass fraction of saturated hydrocarbon in the raw oil, wt%; w2 is the mass fraction of monocyclic aromatic hydrocarbon in the raw oil, wt%; w3 is the mass fraction of aromatic hydrocarbon with more than two rings in the raw oil, wt%; r is a gas constant; t is the reaction temperature of the hydrofining reaction, K;

in the step (2), the conditions of the hydrofinishing reaction include: the reaction temperature is 200 to 300 ℃, the reaction pressure is 1.0 to 4.0MPa, and the liquid hourly space velocity is 2.0 to 10.0 h ^-1 ；

In the step (2), the hydrofining reaction is carried out in the presence of a hydrofining catalyst, the hydrofining catalyst comprises a carrier and a hydrogenation active metal component loaded on the carrier, and the content of the hydrogenation active metal component is 5-50 wt% calculated by oxide based on the total amount of the hydrofining catalyst;

in the step (1), the temperature of the raw oil before being mixed with the hydrogen-containing gas is 200 to 300 ℃;

the volume fraction of hydrogen in the hydrogen-containing gas is 80-99%.

2. The method according to claim 1, wherein, in step (1), the amount of the hydrogen-containing gas to be used is 1~5 times the amount of dissolved hydrogen C in the first mixture, in terms of hydrogen gas contained therein.

3. The process of claim 1, wherein in step (2), the conditions of the hydrofinishing reaction include: the reaction temperature is 220 to 280 ℃, the reaction pressure is 2.0 to 4.0MPa, and the liquid hourly space velocity is 3.0 to 6.0 h ^-1 。

4. The process of claim 1, wherein the hydrogenation active metal component comprises at least one of a group VIB metal element and at least one of a group VIII metal element.

5. The process according to claim 4, wherein the group VIB metal element is molybdenum and/or tungsten and the group VIII metal element is cobalt and/or nickel.

6. The process according to claim 5, wherein the group VIB metal element is present in an amount of from 4 to 40 wt.% in terms of oxides, based on the total amount of the hydrofinishing catalyst; the content of the VIII group metal element is 1 to 10 weight percent calculated by oxide.

7. The process according to claim 6, wherein the group VIB metal element is contained in an amount of 8 to 35 wt.% in terms of oxide, based on the total amount of the hydrofinishing catalyst; the content of the VIII group metal element is 2~5 wt% calculated by oxide.

8. The method according to claim 1 or 2, wherein, in step (1), the introduction pressure of the hydrogen-containing gas is 1.0 to 4.0MPa.

9. The method according to claim 1 or 2, wherein the sulfur content of the raw oil is 200 to 5000 μ g/g; the aromatic hydrocarbon content above the bicyclic ring is not more than 5 wt%.

10. The method according to claim 9, wherein the sulfur content of the raw oil is 200 to 3500 μ g/g; the aromatic hydrocarbon content of the bicyclo ring is not more than 3 wt%.

11. The method of claim 1, wherein the volume fraction of hydrogen in the hydrogen-containing gas is 90% to 99%.

12. The method of claim 1 or 2, wherein the hydrogen-containing gas is from at least one of a hydrogen network of a hydrofinishing unit, a hydrocracking unit, a catalytic reforming unit, and a refinery.