CN103059951B - Catalytic cracking and catalytic gasoline hydrogenation combined technological method - Google Patents

Abstract

The invention discloses a catalytic cracking and catalytic gasoline hydrogenation combined technological method. The method comprises: adjusting the operating conditions of a FCC (fluid catalytic cracking) device fractionating tower, conducting cutting pre-separation on FCC gasoline in the fractionating tower so as to obtain light fractions and heavy fractions; subjecting the light fractions to alkali-free deodorization, then letting the deodorized light fractions and thermocatalytic diesel oil enter a hydrogenation prefractionator together, thus obtaining light gasoline at the tower top and medium gasoline at a tower middle lateral line, mixing the medium gasoline with the heavy fractions, then passing the mixture through a hydrogenation protection reactor and a hydrodesulfurization reactor in order, and mixing the obtained refined product with the deodorized refined light gasoline so as to obtain a clean gasoline product. With the method provided in the invention, the equipment energy consumption is significantly reduced, the gasoline octane number loss is small, gasoline product quality can meet the quality requirement of a sulfur content of less than 10 micrograms/g, and the economic benefits of oil refining enterprises are improved.

Description

A kind of catalytic cracking and gasoline hydrogenation combined technique

Technical field

The present invention relates to a kind of catalytic cracking and gasoline hydrogenation combined technique, is specifically the method that raw material hydrogenation production sulphur content is less than the clean gasoline of 10 μ g/g with catalytic gasoline.

Background technology

Increasingly strict along with environmental regulation, the developed country such as American-European in succession makes laws and proposes more and more stricter regulation to sulphur in motor spirit and olefin(e) centent.From 2009, execution sulphur content is less than 10 μ g/g Europe V emission standards.China requires also more and more stricter to the sulphur content of motor spirit, from 1 day January in 2008, supply Pekinese gasoline starts to perform the specification being equivalent to Europe IV emission standard, and namely sulphur content is less than 50 μ g/g, and similar standard also will be carried out successively in the domestic big city such as Shanghai, Guangzhou.On July 1st, 2010, other areas started to perform the specification being equivalent to EuropeⅢ emission standard, and namely sulphur content is less than 150 μ g/g, and alkene percentage composition is not more than 18v%.As can be seen here, the requirement of following China to content of sulfur in gasoline and olefin(e) centent will be more and more stricter.Therefore, for the product structure of China's motor spirit, be necessary that a kind of new Technology of exploitation is less than the motor spirit of 10 μ g/g for the production of sulphur content, to meet the needs of future market.

Due to historical reasons, in China's motor spirit blend component, catalytically cracked gasoline accounts for about 75% ~ 80%, and has the advantages that sulphur content is higher and alkene is higher.Therefore, reducing China's sulfur content of catalytic cracking gasoline is the major issue faced present stage.

External prior art mainly comprises the SCANFining technique of ExxonMobil company, the Prime-G of Inst Francais Du Petrole ⁺hydrogenating desulfurization/octane value recovering the combination process of technique to be the selective hydrogenation desulfurization process of representative and the OCTGAIN technique of ExxonMobil company, the ISAL technique of Uop Inc. be representative.But because external catalytically cracked gasoline character difference compared with domestic is comparatively large, and proportion is less in gasoline blending component.Therefore, foreign technology is difficult to realize satisfactory results at the domestic catalytically cracked gasoline of processing.

The reducing olefins by hydrogen desulfurization of catalytic gasoline technology of domestic-developed has RSDS, RSDS-II, the RIDOS of Research Institute of Petro-Chemical Engineering and the OCT-M technology of Fushun Petrochemical Research Institute (FRIPP) and OCT-MD technology, these technology all achieve industrialization, but, when production sulphur content is less than the gasoline products of 10 μ g/g, all there is the comparatively large and shortcoming that energy consumption is higher of product loss of octane number.Such as RSDS-II technology of Research Institute of Petro-Chemical Engineering's exploitation, show in the situation of full scale plant running, want the clean gasoline that production sulphur content is less than 10 μ g/g, the loss of octane value will be very large.

CN101307255A discloses a kind of method of producing low sulfur gasoline by using by inferior gasoline fractions.Full cut bad gasoline is first fixed an oxidation deodorizing by the method, mercaptan sulfur is converted into disulphide, then fractionation is lighting end and last running, last running carries out selective hydrodesulfurization through high reactivity/low activity combined hydrogenation desulfurization catalyst, and desulfurization product and lighting end are mixed to get clean gasoline product.Although the method also can produce the gasoline products that sulphur content is less than 10 μ g/g, stock oil adaptability is poor, and loss of octane number is also comparatively large, and technical process and the present invention have very big difference.

CN101787307A discloses a kind of gasoline hydrodesulfurizationmethod method.Gasoline stocks is fractionated into lighting end gasoline and last running gasoline by the method, and wherein the mercaptan sulfur removed wherein refined by lighting end gasoline through alkali cleaning; Last running gasoline, successively through two hydrogenators, carries out hydrogenation and takes off diene, selective hydrodesulfurization and the reaction of selective hydrodesulfurization alcohol; The hydrogenation last running gasoline of gained with refining after lighting end gasoline mix after obtain the full distillation gasoline of super low sulfur.Although the method also can produce the gasoline products that sulphur content is less than 10 μ g/g, raw material has adaptability poor, and technical process is completely different from thinking of the present invention.

Summary of the invention

For the deficiencies in the prior art, the invention provides a kind of catalytic cracking and gasoline hydrogenation combined technique, production sulphur content can be less than clean gasoline or the blend component of 10 μ g/g, and energy consumption significantly reduces compared with existing apparatus.

Catalytic cracking of the present invention and gasoline hydrogenation combined technique comprise following content:

(1) adjust the operation of catalytic cracking unit separation column, in FCC separation column, pre-separation is carried out to FCC gasoline, obtain lighting end and last running; The segmentation temperature of described lighting end and last running is 70 DEG C ~ 85 DEG C;

(2) step (1) gained lighting end enters alkali-free sweetening unit, carries out mercaptan removal process;

(3) lighting end after step (2) gained deodorization enters gasoline hydrogenation preliminary fractionator together with thermocatalysis diesel oil, obtains petroleum naphtha from fractionator overhead, gasoline in lateral line withdrawal function from separation column tower, extracts diesel oil carrying device out at the bottom of tower; The segmentation temperature of petroleum naphtha and middle gasoline is 55 DEG C ~ 70 DEG C; The endpoint control of middle gasoline is 70 DEG C ~ 85 DEG C;

(4), after gasoline mixes with step (1) gained last running in step (3) gained, enter catalytic gasoline pre-hydrogenator, carry out diolefine saturated reaction;

(5) reaction effluent of step (4) enters hydrodesulphurisatioreactors reactors after heat exchange or heat temperature raising, carries out depth-selectiveness hydrogenating desulfurization;

(6) reaction effluent of step (5) enters gas-liquid separator and is separated, and gained product liquid, after air lift, mixes with step (3) gained petroleum naphtha, obtains clean gasoline product or blend component.

According to catalytically cracked gasoline sulfur method of the present invention, catalytic cracking unit described in step (1) comprises various types of catalytic cracking unit, as fluid catalytic cracking (FCC), heavy oil fluid catalytic cracking (RFCC), catalytic pyrolysis (DCC), selective catalysis cracking (SCC), high-yield diesel oil catalytic cracking (MDP), voluminous isomeric olefine catalytic cracking (MIO), voluminous isomeric hydrocarbon catalytic cracking (MIP), voluminous liquefied gas and diesel catalytic cracking (MGD) device etc.

The operation of catalytic cracking unit separation column is adjusted described in step (1), can carry out in newly-built catalytic cracking unit, also can make full use of existing product fractionating system in catalytic cracking unit to transform, such as, can pass through increase catalytic gasoline last running (i.e. heavy petrol) side line and adjust operational condition to realize.Described in step (1), the segmentation temperature of lighting end and last running is generally 70 DEG C ~ 85 DEG C, preferably 75 DEG C ~ 85 DEG C.

Alkali-free sweetening described in step (2) can adopt technology well known in the art.The condition of alkali-free sweetening is generally: reactor operating pressure 0.1MPa ~ 1.0MPa, temperature of reaction 20 DEG C ~ 70 DEG C, Feed space velocities 0.5h ^-1~ 2.0, air flow quantity/inlet amount volume ratio is 0.1 ~ 1.0.Used catalyst and promotor are the catalyzer that this area is commonly used, and can select commercial goods or be prepared according to the knowledge of this area.Lighting end, after alkali-free sweetening, after the mercaptan wherein contained is oxidized to disulphide, enters in heavier middle gasoline.

The feeding manner of the gasoline hydrogenation preliminary fractionator described in step (3) is generally, and lighting end enters from tower bottom, and thermocatalysis diesel oil enters in the middle part of tower.The sideline product that described thermocatalysis diesel oil can be introduced for catalytic cracking main fractionating tower, also can be the catalytic cracking diesel oil introduced before air cooler, doing of diesel oil distillate be generally 330 ~ 380 DEG C.The temperature of thermocatalysis diesel oil is generally 60 DEG C ~ 290 DEG C, preferably 100 DEG C ~ 160 DEG C.In gasoline hydrogenation preliminary fractionator, the segmentation temperature of petroleum naphtha and middle gasoline is 55 DEG C ~ 70 DEG C; Middle gasoline fraction endpoint control is 70 DEG C ~ 85 DEG C, is preferably 75 DEG C ~ 85 DEG C.Lighting end is after alkali-free sweetening, and the mercaptan wherein contained is oxidized to heavier disulphide, enter in the fractionation process of preliminary fractionator heavier in gasoline and diesel oil distillate.

In step (4), after the FCC gasoline last running mixing that middle gasoline and FCC separation column come, as the charging of selective hydrodesulfurization device.The catalyzer that described catalytic gasoline pre-hydrogenator uses is Hydrobon catalyst conventional in this area, as being W-Mo-Ni series hydrocatalyst.The composition of W-Mo-Ni series hydrocatalyst generally includes: Tungsten oxide 99.999 8wt% ~ 15wt%, molybdenum oxide 6wt% ~ 16wt% and nickel oxide 2.0wt% ~ 8.0wt%.The desulfurization catalyst used in hydrodesulphurisatioreactors reactors is also the Hydrobon catalyst that this area is conventional, as being Mo-Co series hydrocatalyst.The composition of Mo-Co series catalysts comprises: molybdenum oxide 6wt% ~ 16 wt%, cobalt oxide 2.0 wt% ~ 8.0 wt%.Hydrogenation products after stripping tower be less than 65 DEG C of light constituents and mix, sulphur content can be obtained lower than the clean gasoline product of 10 μ g/g or blend component.

In the inventive method, in step (4), the operational condition of catalytic gasoline pre-hydrogenator is: hydrogen dividing potential drop 0.8MPa ~ 4.0MPa, best 1.0MPa ~ 2.5MPa; Temperature of reaction is 150 DEG C ~ 250 DEG C, best 160 DEG C ~ 230 DEG C; Volume space velocity is 2.0h ^-1~ 6.0h ^-1, best 2.5h ^-1~ 5.0h ^-1; Hydrogen to oil volume ratio is 10 ~ 300, is preferably 50 ~ 200; The operational condition of hydrodesulphurisatioreactors reactors is: hydrogen dividing potential drop 1.2 MPa ~ 4.0MPa, is preferably 1.5MPa ~ 3.0MPa; Temperature of reaction is 220 DEG C ~ 340 DEG C, preferably at 250 DEG C ~ 320 DEG C; Volume space velocity is 1.0 h ^-1~ 6.0h ^-1, be preferably 2.0 h ^-1~ 4.0 h ^-1; Hydrogen to oil volume ratio is 100 ~ 700, preferably 200 ~ 500.Because two reactors in series use, therefore the working pressure of two reactors is substantially identical, just there is the difference of Pressure Drop; Reaction product is through separator and stripping tower, and product liquid enters product mediation tank field, and the gas circulation being rich in hydrogen returns reactor continuation use.

In step (5), step (4) gained reaction effluent, through process furnace heating or after heating up with the high temperature slurry oil heat exchange of 345 DEG C ~ 500 DEG C that catalytic cracking main fractionating tower is drawn, namely reaches the feeding temperature of hydrodesulphurisatioreactors reactors.Preferably utilize the high temperature slurry oil of catalytic cracking fractionating tower to carry out heat exchange intensification to hydrogenating desulfurization charging, a hydrodesulfurization reaction charging process furnace can be saved like this, save facility investment and running cost.

In step (6), the reaction effluent of step (5) first can carry out heat exchange cooling with the mixture of middle gasoline and last running and hydrogen before entering gas-liquid separator.The hydrogen-rich gas that gas-liquid separator obtains removes after the liquid phase of carrying secretly through cyclone separator and enters desulphurization of recycle hydrogen tower.Desulphurization of recycle hydrogen adopts amine liquid solvent adsorption method, and described amine liquid is organic bases, more with alcamines, conventional Monoethanolamine MEA BASF (MEA), diethanolamine (DEA), diisopropanolamine (DIPA) (DIPA), one or more in N methyldiethanol amine (MDEA).In desulphurization of recycle hydrogen tower, inject poor amine liquid from thionizer top, at the bottom of tower, extract rich amine solution carrying device recycling utilization out; Remove the new hydrogen external with device after compressor boosting of the recycle hydrogen after hydrogen sulfide to mix, as mixed hydrogen for device.Hydrogen sulfide content in described desulphurization of recycle hydrogen Posterior circle hydrogen is 0 ~ 300 μ L/L, preferably 0 ~ 50 μ L/L.

Compared with existing catalyzed gasoline hydrogenation desulfurization technology, the inventive method has following outstanding technique effect:

1, in the inventive method, only need the operational condition adjusting catalytic cracking unit separation column, the pre-separation of the lighting end to catalytic gasoline, last running can be realized in catalytic cracking unit, existing product fractionating system in catalytic cracking unit can be made full use of transform, be conducive to reducing plant modification investment and process cost, and technical scheme mature and reliable, simple.

2, in prior art, oil require that catalytic cracking unit is steamed is after cooling, gasoline hydrogenation device can be transported to by pipeline, and then the requirement of hydrogenation preliminary fractionator feeding temperature just can be reached by heat exchange or process furnace, catalytic gasoline needs the temperature variation that experience cooling, intensification etc. are larger, causes the unnecessary energy consumption of shortening device greatly to increase.And in the inventive method, adopt the flow process of light, the last running direct hot feed of catalytic cracking pre-separation, eliminate the operation of cooling, intensification equitemperature fluctuation, the catalytic gasoline that can make full use of catalytic cracking pre-separation is light, the heat of last running, thus significantly reduce the energy consumption of gasoline hydrogenation preliminary fractionator and hydrogenation unit, reduce facility investment and process cost simultaneously.

3, fractionation is carried out by making the lighting end after deodorization enter hydrogenation preliminary fractionator together with thermocatalysis diesel oil, the heat utilizing the hot diesel oil of high temperature to carry realizes mass transfer, the heat transfer of lighting end and diesel oil distillate, remove the relatively high disulphide of lighting end mid-boiling point after alkali-free sweetening and the easy green coke material of trace, significantly reduce the total sulfur content of tower top petroleum naphtha and middle gasoline.Not only ensure that the mercaptan sulfur content of petroleum naphtha is very low, reduce the content of sulphur content in middle gasoline and coke precursor, extend the running period of pre-hydrogenator, also achieve making full use of of catalytic cracking heat.

4, the charging by utilizing catalytic cracking high temperature slurry oil to heat hydrodesulphurisatioreactors reactors, even can cancel hydrodesulfurization reaction charging process furnace, avoid the coking that hydrogenating desulfurization charging causes in process furnace local heating inequality, the heat of catalytic cracking unit affluence can also be made full use of, reduce the overall energy consumption of gasoline sweetener.

5, the inventive method is for the feature of catalytic gasoline, is divided into different fractions and processes, and while realizing deep desulfuration, reduces product loss of octane number.Catalytic gasoline pre-hydrogenator uses non-precious metal catalyst, at relatively low temperature by saturated for the diolefine in raw material, desulphurization reactor bed coking speed can be slowed down, assurance device running period, because active metallic content is lower, therefore advantage of lower cost.Catalyzer in desulphurization reactor, the while that main purpose being to reduce sulphur content, reduces loss of octane number.

Accompanying drawing explanation

Fig. 1 is the block diagram of catalytic cracking of the present invention and gasoline hydrogenation combined technique.

Embodiment

Below in conjunction with drawings and Examples, the inventive method is done into detailed description.

As shown in Figure 1, in catalytic cracking unit separation column 1, pre-separation is carried out to gasoline, obtain lighting end (comprising petroleum naphtha and middle gasoline) and last running (i.e. heavy petrol).Gained lighting end enters lighting end surge tank 3 through pipeline 2, and enter in alkali-free sweetening reaction tower 5 through feedstock pump 4, lighting end after deodorization enters hydrogenation unit preliminary fractionator 7 with the thermocatalysis diesel oil introduced through pipeline 6, the refining petroleum naphtha obtained at tower top is drawn by pipeline 8, side line obtains middle gasoline and is drawn by pipeline 9, mixes with the incoming stock surge tank 11 of last running through pipeline 10.Mixing raw material oil, through feedstock pump 12, after mixing, enters pre-hydrogenator 13, with catalyst for pre-hydrogenation contact reacts with the hydrogen of pipeline 21.Pre-hydrogenation effluent, after interchanger 14 heat exchange heats up, enters hydrodesulphurisatioreactors reactors 15, with Hydrobon catalyst contact reacts.First hydrogenating desulfurization effluent enters process furnace 14 and heats, and then after interchanger 14 with pre-hydrogenation effluent heat exchange, enters separator 17.Separator 17 gained gas enters desulphurization of recycle hydrogen tower 18 and purifies, and the hydrogen-rich gas after desulfurization enters circulating hydrogen compressor 20 after mixing with the new hydrogen introduced through pipeline 19.Separator 17 gained generates oil and enters stripping tower 22, and gained refined products is drawn through pipeline 23, namely obtains clean gasoline product or blend component after the refining petroleum naphtha of drawing with pipeline 8 is in harmonious proportion.

In the inventive method, the segmentation temperature of the FCC gasoline lighting end that catalytic cracking unit fractionator overhead fractionates out and the FCC gasoline last running of lateral line withdrawal function is 70 DEG C ~ 85 DEG C, preferably 75 DEG C ~ 85 DEG C.Lighting end enters alkali-free sweetening device mercaptan removal, then gasoline hydrogenation device preliminary fractionator is entered together with thermocatalysis diesel oil, tower top obtains petroleum naphtha, in tower, side line obtains middle gasoline, diesel oil is obtained at the bottom of tower, the segmentation temperature of petroleum naphtha and middle gasoline is 55 DEG C ~ 70 DEG C, preferably 60 DEG C ~ 70 DEG C, and doing of middle gasoline is 70 DEG C ~ 85 DEG C.The petroleum naphtha that deodorization aftercut obtains is directly as clean gasoline blend component, after middle gasoline mixes with heavy naphtha, enter gasoline hydrogenation device pre-hydrogenator, the diolefine removed in raw material is contacted with W-Mo-Ni series hydrocatalyst, generation oil enters desulphurization reactor and contacts with Mo-Co series hydrocatalyst, remove the heteroatoms such as sulphur, nitrogen, treated oil mixes with petroleum naphtha after separator and stripping tower, obtains clean gasoline product or blend component.

The catalyzer that catalytic gasoline pre-hydrogenator uses is Hydrobon catalyst conventional in this area, as being W-Mo-Ni series hydrocatalyst.Wherein the composition of W-Mo-Ni series hydrocatalyst comprises: Tungsten oxide 99.999 8wt% ~ 15wt%, molybdenum oxide 6wt% ~ 16wt% and nickel oxide 2.0wt% ~ 8.0wt%.Desulfurization catalyst in hydrodesulphurisatioreactors reactors is also the Hydrobon catalyst that this area is conventional, as being Mo-Co series hydrocatalyst.The composition of catalyzer comprises: molybdenum oxide 6wt% ~ 16wt%, cobalt oxide 2.0wt% ~ 8.0wt%.Support of the catalyst is generally refractory porous oxide, as aluminum oxide, silicon oxide, titanium oxide, zirconium white etc., can contain other adjuvant component.Catalyzer can select existing goods catalyzer, also can prepare by method well known to those skilled in the art.According to the character of feed gasoline, can need load hydrogenation protecting agent on pre-hydrogenator top, Intake Quantity is 5% ~ 20% of pre-hydrogenator hydrogenation catalyst volume, and protectant shape can be Raschig ring, Bird's Nest or abnormal shape etc.

Through the cat naphtha of the inventive method process, product can reach following character: sulphur content is lower than 10 μ g/g, product loss of octane number is less, is less than 1.8 units, and the product after process is applicable to clean gasoline product or blend component as meeting Europe V quality standard.If adopt existing processing method, when production sulphur content is less than the gasoline products of 10 μ g/g, product loss of octane number is comparatively large, and energy consumption is much higher than present method.

In the inventive method, in FCCU separation column, the final boiling point general requirement of gained catalytic gasoline last running is less than 205 DEG C, and sulphur content is less than 600 μ g/g, and olefin(e) centent is less than 28v%.The contents such as concrete technology condition can be determined by those skilled in the art according to material elementses such as the character of raw material, quality product requirements.

The invention has the advantages that:

By catalyzed gasoline hydrogenation desulfurization device is combined with FCC apparatus separation column, not only can ensure that petroleum naphtha removes mercaptan and removes the effect of total sulfur, and the smooth running of complete assembly long period can be ensured, and be conducive to reducing plant modification investment, normal production run comprehensive energy consumption and process cost.

With Conventional catalytic gasoline hydrogenation device process CIMS, after process furnace is changed to desulphurization reactor, first hydrogenating desulfurization effluent is heated, and promoted the temperature in entering hydrodesulphurisatioreactors reactors by heat exchange, can the speed of slowing device coking, prolong operating period.

Hydrogenation preliminary fractionator directly introduced by catalytic unit separation column thermocatalysis diesel oil, for preliminary fractionator provides thermal source, coke precursor in gasoline fraction dissolved simultaneously and takes away, and falls rising to avoid gasoline hydrogenation reactor because of coking build-up of pressure.

The following examples illustrate the present invention further, but be not intended to limit the present invention.

Embodiment 1 ~ 3 adopts the block diagram of Fig. 1, and petroleum naphtha goes out device through pipeline 23 be in harmonious proportion through pipeline 2, hydrogenating desulfurization product, obtains product.The alkali-free sweetening II type technique that alkali-free sweetening adopts Chinese Petroleum Univ. to develop, catalyzer is the AFS-12 prefabrication type catalyzer of Chinese Petroleum Univ.'s research and development.Hydrogenation catalyst is Hydrobon catalyst A and B of industrial application, catalyst A is the FH-40C hydrogenation catalyst of Fushun Petrochemical Research Institute's development and production, catalyst B is the FGH-31 hydrogenation catalyst of Fushun Petrochemical Research Institute's development and production, wherein A is applied to pre-hydrogenator, and B is applied to hydrodesulphurisatioreactors reactors.The character of not carrying out full distillation gasoline during pre-separation is listed in table 1.

Embodiment 1

Carry out pre-separation to catalytic gasoline 1 in catalytic cracking unit separation column, obtain lighting end and last running, the segmentation temperature of described lighting end and last running is 73 DEG C; Alkali-free sweetening is carried out in lighting end, enters hydrogenation preliminary fractionator and carry out fractionation together with the thermocatalysis diesel oil of then drawing with catalytic cracking fractionating tower side line, and the segmentation temperature of gained petroleum naphtha and middle gasoline is 65 DEG C, and doing of middle gasoline is 78 DEG C.Wherein pre-hydrotreating reaction condition is: hydrogen pressure component 1.8MPa, volume space velocity 3.8 h ^-1, temperature of reaction 178 DEG C; Hydrodesulfurization reaction condition is: hydrogen dividing potential drop 1.6MPa, volume space velocity 2.8h ^-1, temperature of reaction 282 DEG C; Total hydrogen to oil volume ratio is 350.

Alkali-free sweetening condition is: reactor operating pressure 0.6MPa, temperature of reaction 35 DEG C, Feed space velocities 0.9h ^-1, air/input material volume ratio is 0.7.Operational condition and test-results list in table 2 and table 3 respectively.

From table 3, adopt this technology that product sulphur content can be made to be down to 10 below μ g/g, product octane value only loses 1.6 units.

Embodiment 2

In catalytic cracking unit separation column, pre-separation is carried out to catalytic gasoline 2, obtain lighting end and last running; The segmentation temperature of described lighting end and last running is 76 DEG C; Alkali-free sweetening is carried out in lighting end, and the thermocatalysis diesel oil of then drawing with catalytic cracking fractionating tower side line enters hydrogenation preliminary fractionator together and is separated, and the segmentation temperature of gained petroleum naphtha and middle gasoline is 62 DEG C, and doing of middle gasoline is 76 DEG C.Wherein pre-hydrotreating reaction condition is: hydrogen pressure component 2.0MPa, volume space velocity 3.5 h ^-1, temperature of reaction 183 DEG C; Hydrodesulfurization reaction condition is: hydrogen dividing potential drop 1.8MPa, volume space velocity 3.0 h ^-1, temperature of reaction 288 DEG C; Total hydrogen to oil volume ratio is 380.

Alkali-free sweetening condition is: reactor operating pressure 0.5MPa, temperature of reaction 45 DEG C, Feed space velocities 0.8h ^-1, air/input material volume ratio is 0.6.Operational condition and test-results list in table 2 and table 3 respectively.

From table 3, adopt this technology that product sulphur content can be made to be down to 10 below μ g/g, product octane value only loses 1.8 units.

Embodiment 3

In catalytic cracking unit separation column, pre-separation is carried out to catalytic gasoline 3, obtain lighting end and last running; The segmentation temperature of described lighting end and last running is 80 DEG C; Alkali-free sweetening is carried out in lighting end, and the thermocatalysis diesel oil of then drawing with catalytic cracking fractionating tower side line enters hydrogenation preliminary fractionator together and is separated, and the segmentation temperature of gained petroleum naphtha and middle gasoline is 67 DEG C, and doing of middle gasoline is 83 DEG C.Wherein pre-hydrotreating reaction condition is: hydrogen pressure component 1.8MPa, volume space velocity 4.2 h ^-1, temperature of reaction 180 DEG C; Hydrodesulfurization reaction condition is: hydrogen dividing potential drop 1.6MPa, volume space velocity 3.2 h ^-1, temperature of reaction 277 DEG C; Total hydrogen to oil volume ratio is 320.

The condition of alkali-free sweetening is: reaction pressure 0.5MPa, temperature of reaction 40 DEG C, Feed space velocities 1.1h ^-1, air/input material volume ratio is 1.0.Operational condition and test-results list in table 2 and table 3 respectively.

From table 3, adopt this technology that product sulphur content can be made to be down to 10 below μ g/g, product octane value only loses 1.2 units.

Comparative example 1

Ordinary method (OCT-MD technology) is adopted to carry out hydrotreatment.Technical process is: technical process is, catalytic cracking full distillate gasoline, after alkali-free sweetening, enters hydrogenation preliminary fractionator and carries out prefractionation, obtains <65 DEG C of lighting end and is greater than the last running of 65 DEG C.Alkali-free sweetening is carried out in <65 DEG C of lighting end, is greater than 65 DEG C of last running and carries out selective hydrodesulfurization.

Stock oil is with embodiment 1, and operational condition and test-results list in table 2 and table 3 respectively.

Table 1 stock oil character

Project	Catalytic gasoline 1	Catalytic gasoline 2	Catalytic gasoline 3	Thermocatalysis diesel oil
					Density, g/cm 3	0.7318	0.7386	0.7238	0.9016
Boiling range, DEG C	42～193	42～196	42～188	298～363
					Sulphur content, μ g/g	421	564	316	9800
Octane value	93.2	92.4	92.8	—
					Temperature, DEG C	—	—	—	296

Table 2 operational condition

Project	Embodiment 1	Embodiment 2	Embodiment 3	Comparative example 1
					Alkali-free sweetening condition
Reaction pressure, MPa	0.6	0.5	0.5	0.6
					Temperature of reaction, DEG C	35	45	40	35
Feed space velocities, h -1	0.9	0.8	1.1	0.9
					Air/input material volume ratio	0.7	0.6	1.0	0.7
Hydroconversion condition
					Reactor	Pre-hydrogenation/desulfurization	Pre-hydrogenation/desulfurization	Pre-hydrogenation/desulfurization	Pre-hydrogenation/desulfurization
Catalyzer	A/B	A/B	A/B	A/B
					Temperature of reaction, DEG C	178/282	183/288	180/277	178/296
Hydrogen dividing potential drop, MPa	1.8/1.6	2.0/1.8	1.8/1.6	1.8/1.6
					Volume space velocity, h -1	3.8/2.8	3.5/3.0	4.2/3.2	3.8/2.8
Hydrogen-oil ratio (always), v/v	350	380	320	350

Table 3 test-results

Project	Embodiment 1	Embodiment 2	Embodiment 3	Comparative example 1
					Density, g/cm 3	0.7309	0.7334	0.7326	0.7313
Boiling range, DEG C, ASTM D86	38～187	38～195	38～187	38～187
					Sulphur content, μ g/g	8.0	9.0	7.6	8.0
RON	91.6	90.6	91.6	90.2
					RON loses	1.6	1.8	1.2	3.0
Plant energy consumption, kgEO/t	9.6	—	—	18
					The cycle of operation, the moon	24～36	24～36	24～36	6~18

As can be seen from Table 3, the sulphur content of product is reduced to 10 below μ g/g, the reaction conditions of the inventive method (embodiment 1) will more relax, and product loss of octane number is only 1.6 units, and ordinary method is 3.0 units simultaneously.In the inventive method, plant energy consumption also drops to 9.6kgEO/t by 18kgEO/t.

Claims (10)

1. catalytic cracking and a gasoline hydrogenation combined technique, comprises following content:

(1) adjust the operation of catalytic cracking unit separation column, in FCC separation column, pre-separation is carried out to FCC gasoline, obtain lighting end and last running; The segmentation temperature of described lighting end and last running is 70 DEG C ~ 85 DEG C;

(2) step (1) gained lighting end enters alkali-free sweetening unit, carries out mercaptan removal process;

(3) lighting end after step (2) gained deodorization enters gasoline hydrogenation preliminary fractionator together with thermocatalysis diesel oil, obtains petroleum naphtha from fractionator overhead, gasoline in lateral line withdrawal function from separation column tower, extracts diesel oil carrying device out at the bottom of tower; The segmentation temperature of petroleum naphtha and middle gasoline is 55 DEG C ~ 70 DEG C; Doing of middle gasoline is 70 DEG C ~ 85 DEG C; The temperature of described thermocatalysis diesel oil is 60 DEG C ~ 290 DEG C;

(4), after gasoline mixes with step (1) gained last running in step (3) gained, enter catalytic gasoline pre-hydrogenator, carry out diolefine saturated reaction;

(5) reaction effluent of step (4) enters hydrodesulphurisatioreactors reactors after heat exchange or heat temperature raising, carries out depth-selectiveness hydrogenating desulfurization; The operational condition of described hydrodesulphurisatioreactors reactors is: hydrogen dividing potential drop 1.2MPa ~ 4.0MPa, and temperature of reaction is 220 DEG C ~ 340 DEG C, and volume space velocity is 1.0h ^-1~ 6.0h ^-1, hydrogen to oil volume ratio is 100 ~ 700;

(6) reaction effluent of step (5) enters gas-liquid separator and is separated, and gained product liquid, after air lift, mixes with step (3) gained petroleum naphtha, obtains clean gasoline product or blend component.

2. in accordance with the method for claim 1, it is characterized in that, described catalytic cracking unit is fluid catalytic cracking, heavy oil fluid catalytic cracking, catalytic pyrolysis, selective catalysis cracking, high-yield diesel oil catalytic cracking, voluminous isomeric olefine catalytic cracking, voluminous isomeric hydrocarbon catalytic cracking or voluminous liquefied gas and diesel catalytic cracking unit.

3. in accordance with the method for claim 1, it is characterized in that, the segmentation temperature of lighting end and last running described in step (1) 75 DEG C ~ 85 DEG C.

4. in accordance with the method for claim 1, it is characterized in that, the condition of the alkali-free sweetening described in step (2) is: reactor operating pressure 0.1 ~ 1.0MPa, temperature of reaction 20 DEG C ~ 70 DEG C, Feed space velocities 0.5 ~ 2.0h ^-1, air flow quantity/inlet amount volume ratio is 0.1 ~ 1.0.

5. in accordance with the method for claim 1, it is characterized in that, the feeding manner of the gasoline hydrogenation preliminary fractionator described in step (3) is, lighting end and thermocatalysis diesel oil enter in the middle part of tower.

6. in accordance with the method for claim 1, it is characterized in that, described thermocatalysis diesel oil is the sideline product that catalytic cracking main fractionating tower is introduced, and doing of diesel oil distillate is 330 ~ 380 DEG C.

7. in accordance with the method for claim 1, it is characterized in that, described in step (4), the operational condition of pre-hydrogenator is: hydrogen dividing potential drop 0.8MPa ~ 4.0MPa, and temperature of reaction is 150 DEG C ~ 250 DEG C, and volume space velocity is 2.0h ^-1~ 6.0h ^-1, hydrogen to oil volume ratio is 10 ~ 300.

8. in accordance with the method for claim 1, it is characterized in that, filling W-Mo-Ni series hydrocatalyst in described pre-hydrogenator, with the weight of catalyzer for benchmark, W-Mo-Ni catalyzer comprises: Tungsten oxide 99.999 8wt% ~ 15wt%, molybdenum oxide 6 wt% ~ 16 wt% and nickel oxide 2.0wt% ~ 8.0wt%.

9. in accordance with the method for claim 1, it is characterized in that, filling Mo-Co series hydrocatalyst in described hydrodesulphurisatioreactors reactors, with the weight of catalyzer for benchmark, catalyzer comprises molybdenum oxide 6 wt% ~ 16 wt%, cobalt oxide 2.0 wt% ~ 8.0 wt%.

10. in accordance with the method for claim 1, it is characterized in that, step (4) gained reaction effluent, after the high temperature slurry oil heat exchange of 345 DEG C ~ 500 DEG C of drawing with catalytic cracking main fractionating tower heats up, reaches the feeding temperature of hydrodesulphurisatioreactors reactors.