CN103059951A - 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 catalytic gasoline hydrogenation combined technique

Technical field

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

Background technology

Increasingly strict along with environmental regulation, the developed country such as American-European in succession makes laws sulphur in the motor spirit and olefin(e) centent has been proposed more and more stricter regulation.From 2009, will carry out sulphur content 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 begins to carry out the specification that is 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 cities such as Shanghai, Guangzhou.On July 1st, 2010, other areas began to carry out the specification that is equivalent to Europe III emission standard, and namely sulphur content is less than 150 μ g/g, and the alkene percentage composition is not more than 18v%.This shows that following China will be more and more stricter to the requirement of content of sulfur in gasoline and olefin(e) centent.Therefore, for the product structure of China's motor spirit, be necessary to develop a kind of new Technology for the production of the motor spirit of sulphur content less than 10 μ g/g, to satisfy the needs of future market.

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

External prior art mainly comprises the SCANFining technique of ExxonMobil company, the Prime-G of Inst Francais Du Petrole ⁺Technique is that the selective hydrogenation desulfurization process of representative and the OCTGAIN technique of ExxonMobil company, the ISAL technique of Uop Inc. are the hydrogenating desulfurization of representative/octane value recovering combination process.But because external catalytically cracked gasoline character is larger with the domestic difference of comparing, and proportion is less in the gasoline blend 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 OCT-M technology and the OCT-MD technology of RSDS, RSDS-II, RIDOS and the Fushun Petrochemical Research Institute (FRIPP) of Research Institute of Petro-Chemical Engineering, these technology have all realized industrialization, but, during less than the gasoline products of 10 μ g/g, all there is the large and higher shortcoming of energy consumption of product loss of octane number in the production sulphur content.The RSDS-II technology of Research Institute of Petro-Chemical Engineering exploitation for example shows that in the situation of full scale plant running want the production sulphur content less than the clean gasoline of 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.The method is fixed an oxidation deodorizing with full cut bad gasoline first, mercaptan sulfur is converted into disulphide, then fractionation is lighting end and last running, last running is carried out selective hydrodesulfurization through high reactivity/low activity combined hydrogenation desulfurization catalyst, and desulfurization product and lighting end are mixed to get the clean gasoline product.Although the method also can be produced sulphur content less than the gasoline products of 10 μ g/g, stock oil adaptability is relatively poor, and loss of octane number is also larger, and technical process and the present invention have very big difference.

CN101787307A discloses a kind of gasoline hydrodesulfurizationmethod method.The method is fractionated into lighting end gasoline and last running gasoline with gasoline stocks, and wherein lighting end gasoline is through the refining mercaptan sulfur that removes wherein of alkali cleaning; Last running gasoline carries out hydrogenation and takes off diene, selective hydrodesulfurization and the reaction of selective hydrodesulfurization alcohol successively through two hydrogenator; The hydrogenation last running gasoline of gained with refining after lighting end gasoline obtain the full distillation gasoline of super low sulfur after mixing.Although the method also can be produced sulphur content less than the gasoline products of 10 μ g/g, raw material has adaptability relatively poor, and technical process is fully 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 catalytic gasoline hydrogenation combined technique, can the production sulphur content less than clean gasoline or the blend component of 10 μ g/g, and energy consumption is compared decrease with existing apparatus.

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

(1) operation of adjustment catalytic cracking unit separation column is carried out pre-separation to FCC gasoline in the FCC separation column, obtains lighting end and last running; The temperature of cutting apart of described lighting end and last running is 70 ℃～85 ℃;

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

(3) lighting end after step (2) the gained deodorization enters catalytic gasoline hydrogenation preliminary fractionator with thermocatalysis diesel oil, obtains petroleum naphtha from fractionator overhead, and gasoline in the lateral line withdrawal function from the fractionation Tata is extracted the diesel oil carrying device out at the bottom of the tower; The temperature of cutting apart of petroleum naphtha and middle gasoline is 55 ℃～70 ℃; The endpoint control of middle gasoline is 70 ℃～85 ℃;

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

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

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

According to catalytically cracked gasoline sulfur method of the present invention, the described catalytic cracking unit of step (1) comprises various types of catalytic cracking unit, such 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 oil catalytic cracking (MGD) device etc.

Adjust the operation of catalytic cracking unit separation column described in the step (1), can carry out in newly-built catalytic cracking unit, also can take full advantage of the interior existing product fractionating system of catalytic cracking unit and transform, for example can realize by increasing catalytic gasoline last running (being a heavy petrol) side line and adjusting operation condition.The temperature of cutting apart of lighting end and last running described in the step (1) is generally 70 ℃～85 ℃, preferred 75 ℃～85 ℃.

Alkali-free sweetening described in the 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, 20 ℃～70 ℃ of temperature of reaction, charging air speed 0.5h ^-1～2.0, air flow quantity/inlet amount volume ratio is 0.1～1.0.Catalyst system therefor and promotor are this area catalyzer commonly used, can select the commercial goods or are prepared according to the knowledge of this area.Lighting end, enters in the heavier middle gasoline after the mercaptan that wherein contains is oxidized to disulphide through behind the alkali-free sweetening.

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

In the step (4), after the FCC gasoline last running that middle gasoline and FCC separation column come mixes, as the charging of selective hydrodesulfurization device.The catalyzer that described catalytic gasoline pre-hydrogenator uses is as Hydrobon catalyst commonly used in this area, as being the 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 that uses in the hydrodesulphurisatioreactors reactors is also as this area Hydrobon catalyst commonly used, as being the 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 through behind the stripping tower with mix less than 65 ℃ of light constituents, can obtain clean gasoline product or blend component that sulphur content is lower than 10 μ g/g.

In the inventive method, the operational condition of catalytic gasoline pre-hydrogenator is in the step (4): hydrogen dividing potential drop 0.8MPa～4.0MPa, preferably 1.0MPa～2.5MPa; Temperature of reaction is 150 ℃～250 ℃, best 160 ℃～230 ℃; Volume space velocity is 2.0h ^-1～6.0h ^-1, best 2.5h ^-1～5.0h ^-1Hydrogen 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 are preferably 1.5MPa～3.0MPa; Temperature of reaction is 220 ℃～340 ℃, is preferably in 250 ℃～320 ℃; Volume space velocity is 1.0 h ^-1～6.0h ^-1, be preferably 2.0 h ^-1～4.0 h ^-1Hydrogen to oil volume ratio is 100～700, is preferably in 200～500.Because two reactors in series are used, therefore the working pressure of two reactors is basic identical, just has 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 that is rich in hydrogen is returned reactor continuation use.

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

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

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

1, in the inventive method, only need adjust the operational condition of catalytic cracking unit separation column, can in catalytic cracking unit, realize the lighting end to catalytic gasoline, the pre-separation of last running, can take full advantage of the interior existing product fractionating system of catalytic cracking unit transforms, be conducive to reduce plant modification investment and process cost, and the technical scheme mature and reliable, simple.

2, in the prior art, oil require that catalytic cracking unit is steamed is through after cooling, can transport to the catalytic gasoline hydrogenation unit by pipeline, and then just can reach the requirement of hydrogenation preliminary fractionator feeding temperature by heat exchange or process furnace, catalytic gasoline need to experience the larger temperature variation such as cooling, intensification, and unnecessary energy consumption increases greatly to cause the shortening device.And adopt the flow process of light, the last running direct hot feed of catalytic cracking pre-separation in the inventive method, cancelled the operation of cooling, intensification equitemperature fluctuation, the catalytic gasoline that can take full advantage of the catalytic cracking pre-separation is light, the heat of last running, thereby the energy consumption of decrease catalytic gasoline hydrogenation preliminary fractionator and hydrogenation unit reduces facility investment and process cost simultaneously.

3, carry out fractionation by making lighting end after the deodorization and thermocatalysis diesel oil enter together the hydrogenation preliminary fractionator, the heat that utilizes the hot diesel oil of high temperature to carry is realized mass transfer, the heat transfer of lighting end and diesel oil distillate, remove easily green coke material of the lighting end mid-boiling point is relatively high behind the alkali-free sweetening disulphide and trace, significantly reduce the total sulfur content of cat head petroleum naphtha and middle gasoline.The mercaptan sulfur content that has not only guaranteed petroleum naphtha is very low, has reduced sulphur content in the middle gasoline and the content of coking precursor, prolongs the running period of pre-hydrogenator, has also realized taking full advantage of of catalytic cracking heat.

4, by utilizing the charging of catalytic cracking high temperature slurry oil heating hydrodesulphurisatioreactors reactors, even can cancel hydrodesulfurization reaction charging process furnace, the coking of having avoided the hydrogenating desulfurization charging to cause in process furnace local heating inequality, can also take full advantage of the heat of catalytic cracking unit affluence, reduce the whole energy consumption of gasoline sweetener.

5, the inventive method is divided into different fractions with it and processes for the characteristics of catalytic gasoline, when realizing deep desulfuration, has reduced the product loss of octane number.The catalytic gasoline pre-hydrogenator uses non-precious metal catalyst, can be under relatively low temperature that the diolefine in the raw material is saturated, slow down desulphurization reactor bed coking speed, assurance device running period, because active metallic content is lower, so cost is relatively low.Catalyzer in the desulphurization reactor, main purpose are when reducing sulphur content, reduce loss of octane number.

Description of drawings

Fig. 1 is the block diagram of catalytic cracking of the present invention and catalytic 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, gasoline is carried out pre-separation in that catalytic cracking unit separation column 1 is interior, obtain lighting end (comprising petroleum naphtha and middle gasoline) and last running (being heavy petrol).The gained lighting end enters lighting end surge tank 3 through pipeline 2, and enter in the alkali-free sweetening reaction tower 5 through feedstock pump 4, lighting end after the deodorization enters hydrogenation unit preliminary fractionator 7 with the thermocatalysis diesel oil of introducing through pipeline 6, drawn by pipeline 8 at the refining petroleum naphtha that cat head obtains, side line obtains middle gasoline and is drawn by pipeline 9, and mixes through the incoming stock surge tank 11 of the last running of pipeline 10.Mixing raw material oil after the hydrogen of pipeline 21 mixes, enters pre-hydrogenator 13, with the catalyst for pre-hydrogenation contact reacts through feedstock pump 12.Pre-hydrogenation effluent enters hydrodesulphurisatioreactors reactors 15, with the Hydrobon catalyst contact reacts after interchanger 14 heat exchange heat up.The hydrogenating desulfurization effluent at first enters process furnace 14 and heats, and then enters separator 17 after interchanger 14 and pre-hydrogenation effluent heat exchange.Separator 17 gained gases enter desulphurization of recycle hydrogen tower 18 and purify, the hydrogen-rich gas after the desulfurization with enter circulating hydrogen compressor 20 after the new hydrogen of introducing through pipeline 19 mixes.Separator 17 gained generate oil and enter stripping tower 22, and the gained refined products is drawn through pipeline 23, namely obtain 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 temperature of cutting apart of the FCC gasoline last running of the FCC gasoline lighting end that the catalytic cracking unit fractionator overhead fractionates out and lateral line withdrawal function is 70 ℃～85 ℃, preferred 75 ℃～85 ℃.Lighting end enters alkali-free sweetening device mercaptan removal, then enter catalytic gasoline hydrogenation unit preliminary fractionator with thermocatalysis diesel oil, cat head obtains petroleum naphtha, side line obtains middle gasoline in the tower, obtain diesel oil at the bottom of the tower, the temperature of cutting apart of petroleum naphtha and middle gasoline is 55 ℃～70 ℃, and preferred 60 ℃～70 ℃, doing of middle gasoline is 70 ℃～85 ℃.The petroleum naphtha that the deodorization aftercut obtains is directly as the clean gasoline blend component, middle gasoline is with after heavy naphtha mixes, enter catalytic gasoline hydrogenation unit pre-hydrogenator, contact the diolefine that removes in the raw material with the W-Mo-Ni series hydrocatalyst, generation oil enters desulphurization reactor and contacts with the Mo-Co series hydrocatalyst, remove the heteroatomss such as sulphur, nitrogen, mix with petroleum naphtha behind treated oil process separator and the stripping tower, obtain clean gasoline product or blend component.

The catalyzer that the catalytic gasoline pre-hydrogenator uses is as Hydrobon catalyst commonly used in this area, as being the 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 the hydrodesulphurisatioreactors reactors is also for this area Hydrobon catalyst commonly used, as being the 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 the refractory porous oxide, such as aluminum oxide, silicon oxide, titanium oxide, zirconium white etc., can contain other adjuvant component.Catalyzer can be selected the existing goods catalyzer, also can be by method preparation well known to those skilled in the art.According to the character of feed gasoline, can need load the hydrogenation protecting agent on pre-hydrogenator top, Intake Quantity is 5%～20% of pre-hydrogenator hydrogenation catalyst volume, protectant shape can be Raschig ring, Bird's Nest or abnormal shape etc.

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

In the inventive method, the final boiling point general requirement of gained catalytic gasoline last running is less than 205 ℃ in the FCCU separation column, 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 raw material properties, quality product requirements.

The invention has the advantages that:

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

Compare with conventional catalytic gasoline hydrogenation unit flow process, after process furnace is changed to desulphurization reactor, first the hydrogenating desulfurization effluent is heated, and promote the temperature in that enters hydrodesulphurisatioreactors reactors by heat exchange, speed that can the slowing device coking, prolong operating period.

Catalytic unit separation column thermocatalysis diesel oil is directly introduced the hydrogenation preliminary fractionator, for preliminary fractionator provides thermal source, simultaneously with coking precursor dissolving in the gasoline fraction and take away, falls rising to avoid the gasoline hydrogenation reactor because of the coking build-up of pressure.

The following examples will the invention will be further described, 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 2, hydrogenating desulfurization product through pipeline 23 and is in harmonious proportion, and obtains product.Alkali-free sweetening adopts the alkali-free sweetening II type technique of Chinese Petroleum Univ.'s exploitation, and catalyzer is the AFS-12 prefabrication type catalyzer of Chinese Petroleum Univ.'s research and development.Hydrogenation catalyst is Hydrobon catalyst A and the 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 full distillation gasoline is not listed in the table 1 when carrying out pre-separation.

Embodiment 1

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

The alkali-free sweetening condition is: reactor operating pressure 0.6MPa, 35 ℃ of temperature of reaction, charging air speed 0.9h ^-1, air/input material volume ratio is 0.7.Operational condition and test-results are listed in respectively table 2 and table 3.

By as seen from Table 3, adopt this technology that the product sulphur content is down to below the 10 μ g/g, the product octane value only loses 1.6 units.

Embodiment 2

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

The alkali-free sweetening condition is: reactor operating pressure 0.5MPa, 45 ℃ of temperature of reaction, charging air speed 0.8h ^-1, air/input material volume ratio is 0.6.Operational condition and test-results are listed in respectively table 2 and table 3.

By as seen from Table 3, adopt this technology that the product sulphur content is down to below the 10 μ g/g, the product octane value only loses 1.8 units.

Embodiment 3

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

The condition of alkali-free sweetening is: reaction pressure 0.5MPa, 40 ℃ of temperature of reaction, charging air speed 1.1h ^-1, air/input material volume ratio is 1.0.Operational condition and test-results are listed in respectively table 2 and table 3.

By as seen from Table 3, adopt this technology that the product sulphur content is down to below the 10 μ g/g, the product octane value only loses 1.2 units.

Comparative Examples 1

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

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

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, ℃	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, ℃	—	—	—	296

Table 2 operational condition

Project	Embodiment 1	Embodiment 2	Embodiment 3	Comparative Examples 1
					The alkali-free sweetening condition	?	?	?	?
Reaction pressure, MPa	0.6	0.5	0.5	0.6
					Temperature of reaction, ℃	35	45	40	35
The charging air speed, 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, ℃	178/282	183/288	180/277	178/296
The 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, ℃, 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
					The RON loss	1.6	1.8	1.2	3.0
Plant energy consumption, kgEO/t	9.6	—	—	18
					The cycle of operation, month	24～36	24～36	24～36	6~18

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

Claims (11)

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

(1) operation of adjustment catalytic cracking unit separation column is carried out pre-separation to FCC gasoline in the FCC separation column, obtains lighting end and last running; The temperature of cutting apart of described lighting end and last running is 70 ℃～85 ℃;

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

(3) lighting end after step (2) the gained deodorization enters catalytic gasoline hydrogenation preliminary fractionator with thermocatalysis diesel oil, obtains petroleum naphtha from fractionator overhead, and gasoline in the lateral line withdrawal function from the fractionation Tata is extracted the diesel oil carrying device out at the bottom of the tower; The temperature of cutting apart of petroleum naphtha and middle gasoline is 55 ℃～70 ℃; Doing of middle gasoline is 70 ℃～85 ℃;

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

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

(6) reaction effluent of step (5) enters gas-liquid separator and separates, and the gained product liquid mixes with step (3) gained petroleum naphtha through after the air lift, 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 oil catalytic cracking unit.

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

4. in accordance with the method for claim 1, it is characterized in that the condition of the alkali-free sweetening described in the step (2) is: reactor operating pressure 0.1～1.0MPa, 20 ℃～70 ℃ of temperature of reaction, charging air speed 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 described catalytic gasoline hydrogenation of step (3) preliminary fractionator is that lighting end and thermocatalysis diesel oil enter from the tower middle part.

6. in accordance with the method for claim 1, it is characterized in that described thermocatalysis diesel oil is the sideline product that the catalytic cracking main fractionating tower is introduced, or the catalytic cracking diesel oil of introducing before the air cooler, doing of diesel oil distillate is 330～380 ℃.

7. in accordance with the method for claim 1, it is characterized in that the temperature of described thermocatalysis diesel oil is 60 ℃～290 ℃.

8. in accordance with the method for claim 1, it is characterized in that the operational condition of hydrogenation protecting reactor is described in the step (4): hydrogen dividing potential drop 0.8MPa～4.0MPa, temperature of reaction is 150 ℃～250 ℃, volume space velocity is 2.0h ^-1～6.0h ^-1, hydrogen to oil volume ratio is 10～300; The operational condition of hydrodesulphurisatioreactors reactors is described in the step (5): hydrogen dividing potential drop 1.2MPa～4.0MPa, and temperature of reaction is 220 ℃～340 ℃, volume space velocity is 1.0h ^-1～6.0h ^-1, hydrogen to oil volume ratio is 100～700.

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

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

11. in accordance with the method for claim 1, it is characterized in that, after step (4) gained reaction effluent passes through 345 ℃ ~ 500 ℃ the high temperature slurry oil heat exchange intensification of drawing with the catalytic cracking main fractionating tower, reach the feeding temperature of hydrodesulphurisatioreactors reactors.