CN103483121B - Method of preparing butadiene from gasoline - Google Patents
Method of preparing butadiene from gasoline Download PDFInfo
- Publication number
- CN103483121B CN103483121B CN201210189757.0A CN201210189757A CN103483121B CN 103483121 B CN103483121 B CN 103483121B CN 201210189757 A CN201210189757 A CN 201210189757A CN 103483121 B CN103483121 B CN 103483121B
- Authority
- CN
- China
- Prior art keywords
- hydrogenation
- weight
- reaction
- gasoline
- hydrogenation reaction
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Active
Links
Landscapes
- Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The method discloses a method of preparing butadiene from gasoline. The method takes a hydrogenation reaction product as at least part of a cracking raw material to carry out steam cracking reactions, wherein the hydrogenation product is prepared according to at least one method picked from following two methods: (1) method one, carrying out hydrogenation reactions with gasoline, whose aromatic hydrocarbon mass content is larger than 12%; (2) method two, separating at least one component of a C6 fraction, a C7 fraction, and a C8 fraction from gasoline, whose aromatic hydrocarbon mass content is larger than 12%, and then subjecting the separated fraction to carry out hydrogenation reactions. The method provided by the invention of preparing butadiene from gasoline has a high yield of butadiene and a low yield of methane, and thus the energy consumption of the cooling box in the separation process of cracking gas is reduced.
Description
Technical field
The present invention relates to a kind of method being prepared divinyl by gasoline.
Background technology
Divinyl is often referred to 1,3-butadiene, also known as ethylene vinyl, is the important basic petrochemical raw material being only second to ethene and propylene.Divinyl is mainly used in synthetic rubber and resin, as polybutadiene rubber (BR), styrene-butadiene rubber(SBR) (SBR), paracril (NBR), styrene-butadiene polymer latex, styrenic thermoplastic elastomer (SBC) and acrylonitrile-butadiene-styrene (ABS) (ABS) resin.In addition, divinyl also can be used for producing the Organic chemical products such as adiponitrile, hexanediamine, nylon-66, BDO, is widely used in numerous areas.
For the preparation of divinyl, traditional method is that the steam crack materials such as lighter hydrocarbons, petroleum naphtha, hydrogenation tail oil are carried out steam cracking.But product main in steam cracking process is ethene, propylene etc., and the yield of divinyl is very low, concrete, the butadiene yield of lighter hydrocarbons steam cracking is about less than 3%, and the butadiene yield of naphtha steam cracking is about 3-6%, and the butadiene yield of hydrogenation tail oil steam cracking is about 4-7%.
The patent applicant of steam cracking process mainly contains U.S. LUMMUS, S & W company of the U.S., KBR company of the U.S., French TECKNIP company, German LINDE company.Steam cracking device is implemented usually in ethylene unit, and ethylene unit mainly comprises pyrolyzer and tripping device.Cracking stock is sent in pyrolyzer and is generated splitting gas through Pintsch process.Splitting gas forms Organic Chemicals and other raw materials, as hydrogen, fuel gas, ethene, propylene, C-4-fraction (comprising butane, butylene, divinyl), pyrolysis gasoline (containing aromatic hydrocarbons), Pyrolysis gas oil PGO, Pyrolysis fuel oil PFO etc. through the isolation andpurification of tripping device.In tripping device, although the processing flow sequence that different patent applicant provides is different, as order separation process, the front-end deethanization flow process of LINDE, the predepropanization process of S & W of LUMMUS, but finally all carry out isolation andpurification according to the carbon number of hydrocarbon.Tripping device comprises oil scrubber, water wash column, ice chest, compressor, demethanizing tower, deethanizing column, ethylene rectification tower, depropanizing tower, propylene rectification tower, debutanizing tower, carbon two and C_3 hydrogenation device, carbon two and carbon three rectifying tower, methanation device and butadiene extraction unit etc.
Splitting gas mainly consists of hydrogen, methane, ethane, ethene, acetylene, propane, propylene, propine and propadiene, butane, butylene, divinyl, carbon five, pyrolysis gasoline (containing aromatic hydrocarbons), Pyrolysis gas oil PGO, Pyrolysis fuel oil PFO, carbon monoxide, carbonic acid gas, butine etc. in addition containing trace.Splitting gas part composition forms raw material, as hydrogen, ethene, propylene, C-4-fraction (comprising butane, butylene, divinyl), pyrolysis gasoline (containing aromatic hydrocarbons) through the separating-purifying of tripping device; Part composition is consumed or recycle, and as carbon monoxide forms fuel gas by methanation device process, methane generates fuel gas by demethanizing tower, and the fuel that fuel gas is used as pyrolyzer is consumed; Carbonic acid gas is absorbed by soda-wash tower; Acetylene, propine and propadiene generate ethene and ethane, propylene and propane through hydrogenator; Ethane, propane form cycle ethane, recycled propane through ethylene rectification tower, propylene rectification tower after purifying, and cycle ethane and recycled propane return pyrolyzer as cracking stock; Pyrolysis gas oil PGO and Pyrolysis fuel oil PFO form oil fuel through oil scrubber.
In splitting gas composition, the economic value added that methane is used as the fuel of pyrolyzer is minimum, and separation of methane needs the splitting gas entering demethanizing tower in ice chest, be cooled to subzero less than 100 DEG C, and energy consumption is very large.Therefore, how reducing the methane yield in splitting gas, is improve one of the economic benefit of ethylene unit, the important channel reducing ethylene unit energy consumption.
In addition, at 1840s, US2391117A proposes under the existence of the hydrocarbon fluid of heat, hexanaphthene is carried out thermo-cracking to prepare the method for divinyl; Propose the method for hexanaphthene being carried out thermo-cracking in US2364377A, GB568536A and GB567913A, need in the method to inject oxygen in reactor, and provide the heat required for reaction by consumption oxygen.But, the butadiene yield of 5-7 % by weight still can only be obtained according to the method for these patent applications.
At present, on the Downstream Market of divinyl, the demand of rubber and ABS resin is continuing to increase, divinyl consumption with average annual about 10% speed increment, and production capacity expansion rate of increase less than 8%, the production of divinyl is still in the state that supply falls short of demand.Therefore, based on steam cracking process, develop new cracking stock and variation route, constantly expand butadiene production approach, for ethylene unit and divinyl industry necessary.
Summary of the invention
The object of the invention is, in order to overcome the defect that in existing butadiene manufacturing process, butadiene yield is lower, to provide a kind of method being prepared divinyl by gasoline.
The invention provides a kind of method being prepared divinyl by gasoline, the method comprises carries out steam cracking reaction using a kind of hydrogenation reaction product as at least part of cracking stock, wherein, described hydrogenation reaction product prepares according at least one in following two kinds of modes:
Mode (1): by aromaticity content be more than 12 % by weight gasoline carry out hydrogenation reaction;
Mode (2): be isolate at least one in C6 cut, C7 cut and C8 cut the gasoline of more than 12 % by weight from aromaticity content, and isolated cut is carried out hydrogenation reaction.
Prepare in the method for divinyl by gasoline described in of the present invention, steam cracking reaction is carried out as at least part of cracking stock by the product using the hydrogenation reaction product of the gasoline of high aromaticity content and/or at least one in C6 cut isolated in the gasoline from high aromaticity content, C7 cut and C8 cut is carried out hydrogenation reaction, can obtain higher butadiene yield, and the yield of methane is lower.
And, because the yield of methane is lower, make the energy consumption in the sepn process of splitting gas relatively low.Particularly, when described method of the present invention is implemented in ethylene unit, in the tripping device of ethylene unit, the energy consumption in the implementation step of ice chest can reduce greatly.
Other features and advantages of the present invention are described in detail in embodiment part subsequently.
Embodiment
Below the specific embodiment of the present invention is described in detail.Should be understood that, embodiment described herein, only for instruction and explanation of the present invention, is not limited to the present invention.
The method preparing divinyl by gasoline according to the present invention comprises carries out steam cracking reaction using a kind of hydrogenation reaction product as at least part of cracking stock, and wherein, described hydrogenation reaction product prepares according at least one in following two kinds of modes:
Mode (1): by aromaticity content be more than 12 % by weight gasoline carry out hydrogenation reaction;
Mode (2): be isolate at least one in C6 cut, C7 cut and C8 cut the gasoline of more than 12 % by weight from aromaticity content, and isolated cut is carried out hydrogenation reaction.
In the present invention, in mode (2), the process being at least one isolated the gasoline of more than 12 % by weight in C6 cut, C7 cut and C8 cut from aromaticity content can be implemented separation method conveniently, such as can by being that the gasoline of more than 12 % by weight carries out distilling realizing by described aromaticity content.Described isolated cut can be C6 cut, C7 cut, C8 cut and their arbitrary combination.
In the preferred case, in order to ensure that method according to the present invention can obtain the butadiene yield significantly improved, in described hydrogenation reaction product, with the total amount of hydrogenation reaction product for benchmark, the total content of hexanaphthene and alkyl cyclohexane is more than 8 % by weight, be more preferably more than 10 % by weight, more preferably more than 20 % by weight, be further preferably more than 30 % by weight.Meanwhile, the total content of described hydrogenation reaction product cyclohexane and alkyl cyclohexane is preferably less than 99 % by weight, is more preferably less than 95 % by weight.
In the present invention, the described hydrogenation reaction in mode (1) and mode (2) all can be implemented under the hydrogenation conditions of routine.Such as, the condition of described hydrogenation reaction can comprise: temperature of reaction is 20-530 DEG C, and reaction pressure is 0.2-50MPa, and hydrogen/oil mol ratio is 0.1-50, and volume space velocity is 0.1-10h
-1.In the present invention, pressure refers to absolute pressure; Described volume space velocity refers to the volume space velocity of the reaction mass not comprising hydrogen, for the hydrogenation process in mode (1), and the volume space velocity of described volume space velocity to be described aromaticity content the be gasoline of more than 12 % by weight; For the hydrogenation process in mode (2), described volume space velocity is the volume space velocity of described isolated cut.
Because described hydrogenation reaction is thermopositive reaction, therefore, for the ease of controlling reaction process, preferred hot(test)-spot temperature in hydrogenator control is 150-250 DEG C further, is preferably 170-230 DEG C.
In the present invention, the process of described hydrogenation reaction can aromatics hydrogenation process conveniently be implemented, such as, can implement according to vapor phase process, liquid phase method, phase transition method or catalytic rectification process.
The implementation process of described vapor phase process can comprise: hydrogen is injected heat exchanger with after the oil phase of hydrogenation mixes, is heated to gas phase; Then, at 200-250 DEG C, the phase feed obtained is injected the first multitubular reactor and reacts, then at 140-180 DEG C, inject the second multitubular reactor react.Wherein, hydrogen and the described mass ratio treating the oil phase of hydrogenation can be 1-10:1, are preferably 3.5-8:1.In reaction process, outside described first reaction tubes and described second reaction tubes, be provided with cooling system, to remove the heat produced in hydrogenation process, reacted in acutely to prevent.
The implementation process of described liquid phase method can comprise: by hydrogen with treat that the oil phase of hydrogenation adds respectively and be equipped with in the main reaction tower of described hydrogenation catalyst, by means of the Circulation of pump, make described hydrogenation catalyst keep suspended state, and remove reaction heat with interchanger, generate low-pressure steam simultaneously; Then, the postreaction tower that described hydrogenation catalyst is housed will be injected again from main reaction tower reaction product out, postreaction tower effluent carries out flash distillation after condensation in high-pressure separation columns, flash gas capable of circulation time main reaction tower, flash liquid injects stabilizer tower, from stabilizer tower tower top removing hydrogen and other gas dissolved, and the logistics of collecting at the bottom of tower is carried out follow-up steam cracking reaction as at least part of cracking stock.The hydrogenation conditions of described liquid phase method preferably includes: temperature of reaction is 150-240 DEG C, and reaction pressure is 0.5-5MPa.
The implementation process of described phase transition method can comprise: by hydrogen with treat that the oil phase of hydrogenation is heated to 130-170 DEG C, then with 1-10h
-1volume space velocity enter bed from the bottom of fixed-bed reactor, carry out liquid-phase hydrogenatin reaction, working pressure is 1-3MPa, and the hot(test)-spot temperature of reactor is 260-280 DEG C, and final reacting product flows out bed from the top of fixed-bed reactor in the form of a vapor.
Described catalytic rectification process is the embodiment that catalyticreactor is coupled outward with rectifying tower, between catalyticreactor and rectifying tower, existing mass-coupling also has thermal coupling, catalyticreactor is preferably fixed-bed tube reactor, reaction liberated heat directly passes to reboiler to produce upflowing vapor by reactor tube walls, a part for rectifying tower tower top phlegma returns in rectifying tower as backflow, to maintain the normal running of tower, another part overhead fraction is with the hydrogen added as reaction raw materials and treat that the oil of hydrogenation mixes, and is fed in catalyticreactor; The mixture of discharging from catalyticreactor gets rid of through gas-liquid separation that non-condensable gas is laggard expects rectifying tower.
Described hydrogenation reaction in mode (1) and mode (2) can carry out hydrogenation in a reactor, also can carry out multistep hydrogenation in a plurality of reactors.
In the present invention, due to described aromaticity content be more than 12 % by weight gasoline and described isolated cut in usually containing diolefin, monoolefine and aromatic hydrocarbons, therefore, in order to prevent occurring temperature runaway phenomenon in described hydrogenation process, described hydrogenation process in mode (1) and mode (2) preferably comprises the first hydrogenation process separately successively, second hydrogenation process and the 3rd hydrogenation process, described first hydrogenation process is carried out under diolefin hydrogenation conditions, described second hydrogenation process is carried out under monoolefin hydrogenation reaction conditions, described 3rd hydrogenation process is carried out under aromatic hydrogenation reaction conditions.Described first hydrogenation process, also referred to as diolefin hydrogenation process, is mainly used in diolefin component to be converted into monoolefine.Described second hydrogenation process, also referred to as monoolefin hydrogenation process, is mainly used in monoolefine component to be converted into alkane.Described 3rd hydrogenation process is also referred to as aromatic hydrogenation process, and being mainly used in aromatic conversion is naphthenic hydrocarbon.
In the present invention, all there is no particular limitation separately for described diolefin hydrogenation conditions, described monoolefin hydrogenation reaction conditions and described aromatic hydrogenation reaction conditions, suitably can select respectively in the aromatic hydrogenation reaction conditions of the diolefin hydrogenation conditions of routine, conventional monoolefin hydrogenation reaction conditions and routine.Under preferable case, described diolefin hydrogenation conditions comprises: temperature of reaction is 30-250 DEG C, and reaction pressure is 1-10MPa, and hydrogen/oil mol ratio is 0.1-20, and volume space velocity is 0.1-10h
-1; Described monoolefin hydrogenation reaction conditions comprises: temperature of reaction is 240-400 DEG C, and reaction pressure is 1-10MPa, and hydrogen/oil mol ratio is 0.5-50, and volume space velocity is 0.1-10h
-1; Described aromatic hydrogenation reaction conditions comprises: temperature of reaction is 20-250 DEG C, and reaction pressure is 0.2-50MPa, and hydrogen/oil mol ratio is 0.1-10, and volume space velocity is 0.1-10h
-1.
In described hydrogenation process, in order to prevent the heat accumulated in hydrogenation reaction system too much, preferably also comprise in the present invention and hydrogenation reaction product or hexanaphthene and/or alkyl cyclohexane are diluted the material in hydrogenation reaction system, or cooling system (as water cooling system) is installed in hydrogenation reaction device, to take away the heat produced in hydrogenation process.
In the present invention, the hydrogenation reaction in mode (1) and mode (2) is all carried out under the existence of hydrogenation catalyst.For described hydrogenation catalyst, the various hydrogenation catalysts that this area routine uses all are applicable to the present invention.Usually, described hydrogenation catalyst can be carried on supported catalyst for hydrogenation component.Under preferable case, described hydrogenation catalyst is the hydrogenation catalyst comprising group VIII metal and/or group vib metal, and that is, described hydrogenation component is preferably group VIII metal and/or group vib metal.Described group VIII metal can be such as platinum (Pt), palladium (Pd), nickel (Ni), ruthenium (Ru), rhodium (Rh) or iridium (Ir).Described group vib metal can be such as molybdenum (Mo) or tungsten (W).As described hydrogenation component, group VIII metal and group vib the metal separately usual form with corresponding metal oxide and/or metal halide exist.
Described carrier can be the various support of the catalyst that this area routine uses.Described carrier can be heat-resistant inorganic oxide carrier, such as aluminum oxide.
In described hydrogenation catalyst, the content of described hydrogenation component can be 0.1-50 % by weight, is preferably 1-40 % by weight; The content of described carrier can be 50-99.9 % by weight, is preferably 60-99 % by weight.
For described hydrogenation catalyst, suitably select in hydrogenation catalyst product that can be commercially available.Concrete, described hydrogenation catalyst can be such as palladium series catalyst, as SHP catalyst series (such as SHP-01F and SHP-01S) and LY catalyst series (such as LY-9801 and LY-9802); Nickel catalyst, as HTC catalyst series (such as HTC-200 and HTC-400), HC-402-2 type Ni-based homogeneous phase Ziegler type complex catalyst and NCG-6 type benzene hydrogenating catalyst; Platinum group catalyst, as NCH1-1 type benzene hydrogenating catalyst and MCH-1 type hydrogenation catalyst.
In the preferred case, when the described hydrogenation process in mode (1) and mode (2) comprises above-mentioned three hydrogenation processes (i.e. diolefin hydrogenation process, monoolefin hydrogenation process and aromatic hydrogenation process) separately, for described diolefin hydrogenation process, the hydrogenation catalyst adopted is palladium series catalyst and/or nickel catalyst, described palladium series catalyst refers to that active ingredient comprises the hydrogenation catalyst of palladium, such as, can be commercially available SHP catalyst series (such as SHP-01F and SHP-01S) and LY catalyst series (such as LY-9801); Described nickel catalyst refers to that active ingredient comprises the hydrogenation catalyst of nickel, such as, can be commercially available HTC catalyst series (such as HTC-200 and HTC-400).For described monoolefin hydrogenation process, the hydrogenation catalyst adopted is Co-Mo-Ni catalyst series, described Co-Mo-Ni catalyst series refers to that active ingredient comprises the hydrogenation catalyst of cobalt, molybdenum and nickel three, such as, can be commercially available SHP catalyst series (such as SHP-02 and SHP-02F), LY catalyst series (such as LY-9802) and BY catalyst series (as BY-5).For described aromatic hydrogenation process, the hydrogenation catalyst adopted is nickel catalyst and/or platinum group catalyst, described nickel catalyst can be such as commercially available DYX-8 nickel-alumina catalyst and NCG catalyst series (such as NCG-98H type catalyzer), described platinum group catalyst refers to that active ingredient comprises the hydrogenation catalyst of platinum, such as, can be commercially available NCH1-1 type benzene hydrogenating catalyst and MCH-1 type hydrogenation catalyst.
In the present invention, the aromaticity content in described gasoline is more than 12 % by weight.But, usual aromaticity content be more than 12 % by weight cracking stock easily cause the waste heat boiler coking of industrial pyrolysis furnace serious, the cycle of operation of very big shortening industrial pyrolysis furnace, thus aromaticity content be more than 12 % by weight hydrocarbon feed directly can not be used as cracking stock.Therefore, employing aromaticity content of the present invention is the gasoline method of preparing divinyl of more than 12 % by weight to be aromaticity content be more than 12 % by weight gasoline provide a kind of new utilization ways.In the preferred case, the aromaticity content in described gasoline is 12-95 % by weight.
In the present invention, the hydrogenation reaction product that mode (1) and/or mode (2) produce can be used alone as cracking stock, also can mix with the cracking stock of routine and be used as cracking stock.The cracking stock of described routine can be such as lighter hydrocarbons, petroleum naphtha, diesel oil, hydrogenation tail oil etc.
In the present invention, described steam cracking reaction can carry out in pyrolyzer.In the preferred case, in order to prevent the convection zone coking of pyrolyzer, preferably by used in combination for the cracking stock of described hydrogenation reaction product and described routine.Preferred, by described hydrogenation reaction product and petroleum naphtha used in combination, also namely in described cracking stock, except described hydrogenation reaction product, also containing petroleum naphtha.Further preferred, described steam crack material contains the petroleum naphtha of 1-80 % by weight and the described hydrogenation reaction product of 20-99 % by weight, more preferably containing the petroleum naphtha of 1-70 % by weight and the described hydrogenation reaction product of 30-99 % by weight.
Described method of the present invention is preferably implemented in ethylene unit.Described ethylene unit comprises pyrolyzer and tripping device.Described pyrolyzer can be the pyrolyzer of the preparing ethylene by steam cracking that this area routine uses.Described pyrolyzer mainly comprises convection zone, radiation section, quenching boiler and gas burning system usually.In described pyrolyzer, by cracking stock be steam heated to generation steam cracking reaction, generate and be rich in the splitting gas of ethene.In the preferred case, described pyrolyzer is preferably tube cracking furnace.Described tube cracking furnace comprises convection zone, radiation section, quenching boiler and gas burning system, and cracking stock enters radiation section in convection zone; In radiation section, cracking stock be steam heated to generation steam cracking reaction, generate and be rich in the splitting gas of ethene; Splitting gas out enters quenching boiler afterwards from radiation section, and in quenching boiler, splitting gas is cooled to 300-600 DEG C, to make splitting gas scission reaction not occur, reclaims heat simultaneously; Fuel system is used for providing heat to steam cracking reaction process.Described tripping device is used for hydrocarbon splitting gas being separated into different carbon number.Usually, described tripping device mainly comprises: oil scrubber, water wash column, ice chest, compressor, demethanizing tower, deethanizing column, ethylene rectification tower, depropanizing tower, propylene rectification tower, debutanizing tower, carbon two and C_3 hydrogenation device, carbon two and carbon three rectifying tower, methanation device and butadiene extraction unit.The implementation method of described tripping device has been conventionally known to one of skill in the art, does not repeat them here.In the present invention, because the methane yield in steam-cracking process is lower, the methane content in splitting gas is lower, and the energy consumption in thus sepn process (particularly, the implementation step of ice chest) is lower.
In the present invention, when described steam cracking reaction carries out in pyrolyzer, in described steam cracking reaction process, the coil outlet temperature of described pyrolyzer is preferably 760-890 DEG C, is more preferably 810-840 DEG C; Water weight of oil, than being 0.3-1, is preferably 0.45-0.65.In addition, in described steam cracking reaction process, other Parameter Conditions of described pyrolyzer can processing condition conveniently be implemented, and there is no particular limitation in the present invention.
The invention will be further described by the following examples, but protection scope of the present invention is not limited in these embodiments.
In the following Examples and Comparative Examples, butadiene yield and methane yield are according to following formulae discovery:
Gross weight × 100% of the weight/steam cracking reaction product of divinyl in butadiene yield=steam cracking reaction product
Gross weight × 100% of the weight/steam cracking reaction product of methane in methane yield=steam cracking reaction product.
For the petroleum naphtha used in following examples and comparative example, shown in the correlation parameter of this petroleum naphtha and the table 1 composed as follows that records according to ASTM D5443 method.
Table 1
Embodiment 1
The present embodiment is for illustration of the described method being prepared divinyl by gasoline of the present invention.
(1) hydrogenation
Cracking stock is used as to after pyrolysis gasoline (physical parameter is as shown in table 2 below) direct hydrogenation, hydrogenation process is specific as follows: carry out in four hydrogenators pyrolysis gasoline hydrogenation, first is diolefin hydrogenator, second is monoolefin hydrogenation reactor, and third and fourth is the aromatics hydrogenation reactors of connecting.In diolefin hydrogenator, the catalyzer of employing be LY-9801(purchased from Lanzhou Petrochemical Company chemical research institute, Pd content is 0.3 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 60 DEG C, and absolute pressure is 2.6MPa, and hydrogen/oil mol ratio is 0.6, and the volume space velocity of pyrolysis gasoline is 3h
-1.In monoolefin hydrogenation reactor, the catalyzer of employing be LY-9802(purchased from Lanzhou Petrochemical Company chemical research institute, Co content is 9.35 % by weight, Mo content be 17.28 % by weight, Ni content is 7.11 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 280 DEG C, and absolute pressure is 5.67MPa, and hydrogen/oil mol ratio is 4, and the volume space velocity from the material of diolefin hydrogenator is 2.5h
-1.In two aromatic hydrocarbons serial hydrogenation reactors, the catalyzer of employing is NCG-98H type catalyzer, and (purchased from Nanjing Chuan Ming catalyst technology limited liability company, NiO content is 39.29 % by weight, and carrier is Al
2o
3), hydrogen/oil mol ratio is 7, and front aromatics hydrogenation reactors temperature in is 150 DEG C, and hot(test)-spot temperature is 170-220 DEG C, and absolute pressure is 0.5MPa, and the volume space velocity from the material of monoolefin hydrogenation reactor is 1.0h
-1.Learn by analysis, in the hydrogenation reaction product obtained after the aromatics hydrogenation reactors through series connection, cyclohexane content is 34.51 % by weight, and methylcyclohexane is 10.65 % by weight, and dimethyl cyclohexane is 3.45 % by weight.
Table 2
| Composition | Content (% by weight) |
| Normal paraffin | 10.12 |
| Isoparaffin | 11.83 |
| Chain monoene | 4.3 |
| Chain diene | 11.39 |
| Cyclopentenes | 0.8 |
| Cyclopentadiene | 8.09 |
| Cyclohexadiene | 3.11 |
| Tetrahydrobenzene | 0.37 |
| Hexanaphthene | 0.9 |
| Benzene | 30.28 |
| Methyl ring two hexene | 0.1 |
| Tetrahydrotoluene | 0.36 |
| Methylcyclohexane | 0.66 |
| Toluene | 9.58 |
| Ethyl-cyclohexene | 0.23 |
| Ethylcyclohexane | 0.65 |
| Dimethylbenzene | 2.58 |
| Indenes | 1.09 |
| Carbon 9 and above aromatic hydrocarbons | 3.56 |
(2) steam cracking
The hydrogenation reaction product obtained after the aromatics hydrogenation reactors of series connection in step (1) is injected CBL-III type pyrolyzer (purchased from China PetroChemical Corporation) and carries out steam cracking reaction, wherein, inlet amount is 24.076 tons/hour, water weight of oil ratio is 0.5, and coil outlet temperature is 840 DEG C.Collect splitting gas (i.e. steam cracking reaction product), analyze composition wherein, and calculate butadiene yield and methane yield, result is as shown in table 6 below.Then, in the tripping device be made up of with butadiene extraction unit oil scrubber, water wash column, ice chest, compressor, demethanizing tower, deethanizing column, ethylene rectification tower, depropanizing tower, propylene rectification tower, debutanizing tower, carbon two and C_3 hydrogenation device, carbon two and carbon three rectifying tower, methanation device successively, splitting gas is separated, and recording the energy consumption of ice chest, result is as shown in table 6 below.
Comparative example 1
Method according to embodiment 1 carries out steam cracking reaction, and difference is, the material injecting CBL-III type pyrolyzer is petroleum naphtha.
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Comparative example 2
It is in the molten salt bath of 761 DEG C that snake pipe reactor is immersed temperature, by in propane flow in pipes reactor, simultaneously by hexanaphthene vaporization also flow in pipes reactor, with the stream contacts of propane, there is heat scission reaction, and the ratio of distance between the main-inlet of distance between the inlet of hexanaphthene of vaporization and the outlet of pipeline reactor and propane and the outlet of pipeline reactor is 2:3, the volume ratio of the injection rate of the injection rate of propane and the hexanaphthene of vaporization is 3:1, the hexanaphthene of vaporization is identical with the flow velocity of propane, and the hexanaphthene that can calculate vaporization according to flow velocity and propane are contacting the outlet reaching pipeline reactor after about 0.6 second.Collect splitting gas (instant heating cleavage reaction product) from the exit of pipeline reactor, analyze composition wherein, and calculate butadiene yield and methane yield, result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Comparative example 3
The tubular reactor being filled with 8-14 object quartz chips is heated to 750 DEG C, by the hexanaphthene containing 0.9 % by mole of oxygen with 5530h
-1air speed inject tubular reactor carry out heat scission reaction.Collect splitting gas (instant heating cleavage reaction product), analyze composition wherein, and calculate butadiene yield and methane yield, result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Embodiment 2
The present embodiment is for illustration of the described method being prepared divinyl by gasoline of the present invention.
Method according to embodiment 1 carries out hydrogenation and steam cracking, difference is, in steam-cracking process, hydrogenation reaction product is injected CBL-III type pyrolyzer (purchased from China PetroChemical Corporation) with petroleum naphtha with the mixture that the weight ratio of 1:1 mixes and carries out steam cracking reaction.
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Embodiment 3
The present embodiment is for illustration of the described method being prepared divinyl by gasoline provided by the invention.
Method according to embodiment 1 carries out hydrogenation and steam cracking, difference is, in steam-cracking process, hydrogenation reaction product is injected CBL-III type pyrolyzer (purchased from China PetroChemical Corporation) with petroleum naphtha with the mixture that the weight ratio of 1:4 mixes and carries out steam cracking reaction.
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Embodiment 4
The present embodiment is for illustration of the described method being prepared divinyl by gasoline of the present invention.
(1) hydrogenation
Pyrolysis gasoline (physical parameter is as above shown in table 2) is distilled, collect the C6-C8 cut that boiling point is 65-167 DEG C, shown in its table 3 composed as follows, then hydrogenation is carried out to this C6-C8 cut, hydrogenation process is specific as follows: carry out in four hydrogenators C6-C8 cut fraction hydrogenation, first is diolefin hydrogenator, and second is monoolefin hydrogenation reactor, and third and fourth is the aromatics hydrogenation reactors of connecting.In diolefin hydrogenator, the catalyzer of employing be SHP-01F(purchased from Shang Petrochemical Inst., SINOPEC, Pd content is 0.4 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 150 DEG C, and absolute pressure is 10MPa, and hydrogen/oil mol ratio is the volume space velocity of 5, C6-C8 cut is 6h
-1.In monoolefin hydrogenation reactor, the catalyzer of employing be SHP-02(purchased from Shang Petrochemical Inst., SINOPEC, Co content is 2.75 % by weight, Mo content be 10.33 % by weight, Ni content is 2.76 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 400 DEG C, and absolute pressure is 10MPa, and hydrogen/oil mol ratio is 15, and the volume space velocity from the material of diolefin hydrogenator is 6h
-1.In two aromatic hydrocarbons serial hydrogenation reactors, the catalyzer of employing is DYX-8 nickel-alumina catalyst, and (purchased from grand celebration Petroleum Institute, Ni content is 17.65 % by weight, and carrier is Al
2o
3), hydrogen/oil mol ratio is 10, and front aromatics hydrogenation reactors temperature in is 200 DEG C, and hot(test)-spot temperature is 170-220 DEG C, and absolute pressure is 10MPa, and the volume space velocity from the material of monoolefin hydrogenation reactor is 3h
-1.Learn by analysis, in the hydrogenation reaction product obtained after the aromatics hydrogenation reactors through series connection, cyclohexane content is 59.77 % by weight, and methylcyclohexane is 18.45 % by weight, and dimethyl cyclohexane is 5.97 % by weight.
Table 3
| Composition | Content (% by weight) |
| Benzene | 52.44 |
| Chain carbon 6 diolefine | 1.40 |
| Chain carbon 6 monoolefine | 1.61 |
| Normal hexane | 3.38 |
| Chain carbon 6 isoparaffin | 2.98 |
| Cyclohexadiene | 5.39 |
| Tetrahydrobenzene | 0.64 |
| Hexanaphthene | 1.56 |
| Toluene | 16.59 |
| Chain carbon 7 diolefine | 2.46 |
| Chain carbon 7 monoolefine | 0.88 |
| Normal heptane | 0.97 |
| Chain carbon 7 isoparaffin | 0.80 |
| Methyl cyclohexadiene | 0.17 |
| Tetrahydrotoluene | 0.62 |
| Methylcyclohexane | 1.14 |
| Ethylbenzene | 1.49 |
| M-xylene | 2.27 |
| P-Xylol | 0.71 |
| Chain carbon 8 alkene | 0.24 |
| Octane | 0.28 |
| Chain carbon 8 isoparaffin | 0.45 |
| Ethyl-cyclohexene | 0.40 |
| Ethylcyclohexane | 1.13 |
(2) steam cracking
The hydrogenation reaction product obtained after the aromatics hydrogenation reactors of series connection in step (1) is injected CBL-III type pyrolyzer (purchased from China PetroChemical Corporation) and carries out steam cracking reaction, wherein, inlet amount is 24.076 tons/hour, water weight of oil ratio is 0.5, and coil outlet temperature is 840 DEG C.
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Embodiment 5
The present embodiment is for illustration of the described method being prepared divinyl by gasoline of the present invention.
(1) hydrogenation
Pyrolysis gasoline (physical parameter is as above shown in table 2) is distilled, collect the C6-C7 cut that boiling point is 65-120 DEG C, shown in its table 4 composed as follows, then hydrogenation is carried out to this C6-C7 cut, hydrogenation process is specific as follows: carry out in four hydrogenators C6-C7 cut fraction hydrogenation, first is diolefin hydrogenator, and second is monoolefin hydrogenation reactor, and third and fourth is the aromatics hydrogenation reactors of connecting.In diolefin hydrogenator, the catalyzer of employing be HTC-200(purchased from Synetix company of Britain, Ni content is 16.5 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 30 DEG C, and absolute pressure is 1MPa, and hydrogen/oil mol ratio is the volume space velocity of 10, C6-C7 cut is 1h
-1.In monoolefin hydrogenation reactor, the catalyzer of employing be SHP-02F(purchased from Shang Petrochemical Inst., SINOPEC, Mo content is 11.33 % by weight, Ni content is 3.14 % by weight, and carrier is Al
2o
3), reactor inlet temperature is 240 DEG C, and absolute pressure is 1MPa, and hydrogen/oil mol ratio is 0.5, and the volume space velocity from the material of diolefin hydrogenator is 0.5h
-1.In two aromatic hydrocarbons serial hydrogenation reactors, the catalyzer of employing is NCG-98H catalyzer, and (purchased from Nanjing Chuan Ming catalyst technology limited liability company, NiO content is 39.29 % by weight, and carrier is Al
2o
3), hydrogen/oil mol ratio is 6, and front aromatics hydrogenation reactors temperature in is 50 DEG C, and hot(test)-spot temperature is 170-220 DEG C, and absolute pressure is 2MPa, and the volume space velocity from the material of monoolefin hydrogenation reactor is 6h
-1.Learn by analysis, in the hydrogenation reaction product obtained after the aromatics hydrogenation reactors through series connection, cyclohexane content is 64.24 % by weight, and methylcyclohexane is 19.83 % by weight.
Table 4
| Composition | Content (% by weight) |
| Benzene | 56.37 |
| Chain carbon 6 diolefine | 1.51 |
| Chain carbon 6 monoolefine | 1.73 |
| Normal hexane | 3.63 |
| Chain carbon 6 isoparaffin | 3.20 |
| Cyclohexadiene | 5.79 |
| Tetrahydrobenzene | 0.69 |
| Hexanaphthene | 1.68 |
| Toluene | 17.83 |
| Chain carbon 7 diolefine | 2.64 |
| Chain carbon 7 monoolefine | 0.95 |
| Normal heptane | 1.04 |
| Chain carbon 7 isoparaffin | 0.86 |
| Methyl cyclohexadiene | 0.19 |
| Tetrahydrotoluene | 0.67 |
| Methylcyclohexane | 1.23 |
(2) steam cracking
The hydrogenation reaction product obtained after the aromatics hydrogenation reactors of series connection in step (1) is injected CBL-III type pyrolyzer (purchased from China PetroChemical Corporation) and carries out steam cracking reaction, wherein, inlet amount is 24.076 tons/hour, water weight of oil ratio is 0.5, and coil outlet temperature is 810 DEG C.
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Embodiment 6
The present embodiment is for illustration of the described method being prepared divinyl by gasoline of the present invention.
Method according to embodiment 1 carries out hydrogenation and steam cracking, and difference is, the reaction raw materials of hydrogenation process is: distill pyrolysis gasoline (physical parameter is as above shown in table 2), collect the C6 cut that boiling point is 65-85 DEG C, shown in its table 5 composed as follows.
Table 5
| Composition | Content (% by weight) |
| Benzene | 75.57 |
| Chain carbon 6 diolefine | 2.02 |
| Chain carbon 6 monoolefine | 2.32 |
| Normal hexane | 4.87 |
| Chain carbon 6 isoparaffin | 4.29 |
| Cyclohexadiene | 7.76 |
| Tetrahydrobenzene | 0.92 |
| Hexanaphthene | 2.25 |
Method according to embodiment 1 is analyzed splitting gas, and calculates butadiene yield and methane yield, and result is as shown in table 6 below.Then, the method according to embodiment 1 is separated splitting gas, and records the energy consumption of ice chest, and result is as shown in table 6 below.
Table 6
As can be seen from the data of table 6, the method preparing divinyl by gasoline according to the present invention can obtain the butadiene yield and lower methane yield that significantly improve, and reduces the energy consumption in sepn process.Concrete, by embodiment 1 and comparative example 1 are compared and can find out, the cracking stock used in comparative example 1 is petroleum naphtha, and the hydrogenation reaction product that the cracking stock used in embodiment 1 is the pyrolysis gasoline of high aromaticity content, acetonideexample 1 obtains the butadiene yield significantly improved, and ice chest energy consumption in splitting gas sepn process is obviously lower; Compare by embodiment 1 can find out with comparative example 2-3, that hexanaphthene is carried out thermo-cracking in comparative example 2 and 3, and be the hydrogenation reaction product of the pyrolysis gasoline of high aromaticity content (mainly hexanaphthene and alkyl cyclohexane) is carried out steam cracking in embodiment 1, acetonideexample 1 obtains the methane yield of butadiene yield and the reduction significantly improved, and reduces the ice chest energy consumption in splitting gas sepn process.
Claims (10)
1. prepared a method for divinyl by gasoline, the method comprises carries out steam cracking reaction using a kind of hydrogenation reaction product as at least part of cracking stock, and wherein, described hydrogenation reaction product prepares according at least one in following two kinds of modes:
Mode (1): by aromaticity content be more than 12 % by weight gasoline carry out hydrogenation reaction;
Mode (2): be isolate at least one in C6 cut, C7 cut and C8 cut the gasoline of more than 12 % by weight from aromaticity content, and isolated cut is carried out hydrogenation reaction,
Described hydrogenation process in mode (1) and mode (2) comprises the first hydrogenation process, the second hydrogenation process and the 3rd hydrogenation process separately successively, described first hydrogenation process is carried out under diolefin hydrogenation conditions, described second hydrogenation process is carried out under monoolefin hydrogenation reaction conditions, described 3rd hydrogenation process is carried out under aromatic hydrogenation reaction conditions
Wherein, described diolefin hydrogenation conditions comprises: temperature of reaction is 30-250 DEG C, and reaction pressure is 1-10MPa, and hydrogen/oil mol ratio is 0.1-20, and volume space velocity is 0.1-10h
-1; Described monoolefin hydrogenation reaction conditions comprises: temperature of reaction is 240-400 DEG C, and reaction pressure is 1-10MPa, and hydrogen/oil mol ratio is 0.5-50, and volume space velocity is 0.1-10h
-1; Described aromatic hydrogenation reaction conditions comprises: temperature of reaction is 20-250 DEG C, and reaction pressure is 0.2-50MPa, and hydrogen/oil mol ratio is 0.1-10, and volume space velocity is 0.1-10h
-1.
2. method according to claim 1, wherein, in described hydrogenation reaction product, with the total amount of hydrogenation reaction product for benchmark, the total content of hexanaphthene and alkyl cyclohexane is more than 8 % by weight.
3. method according to claim 1 and 2, wherein, in described hydrogenation reaction product, with the total amount of hydrogenation reaction product for benchmark, the total content of hexanaphthene and alkyl cyclohexane is more than 30 % by weight.
4. method according to claim 1 and 2, wherein, hydrogenation reaction in mode (1) and mode (2) is all carried out under the existence of hydrogenation catalyst, and described hydrogenation catalyst is separately for comprising the hydrogenation catalyst of group VIII metal and/or group vib metal.
5. method according to claim 1, wherein, the aromaticity content in described gasoline is 12-95 % by weight.
6. method according to claim 1, wherein, except described hydrogenation reaction product, described cracking stock is also containing petroleum naphtha.
7. method according to claim 6, wherein, in described cracking stock, the content of petroleum naphtha is 1-80 % by weight; The content of described hydrogenation reaction product is 20-99 % by weight.
8. method according to claim 7, wherein, in described cracking stock, the content of petroleum naphtha is 1-70 % by weight; The content of described hydrogenation reaction product is 30-99 % by weight.
9. according to the method in claim 1 and 6-8 described in any one, wherein, described steam cracking reaction carries out in pyrolyzer, and in described steam cracking reaction process, the coil outlet temperature of described pyrolyzer is 760-890 DEG C; Water weight of oil is than being 0.3-1.
10. method according to claim 9, wherein, the coil outlet temperature of described pyrolyzer is 810-840 DEG C; Water weight of oil is than being 0.45-0.65.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CN201210189757.0A CN103483121B (en) | 2012-06-08 | 2012-06-08 | Method of preparing butadiene from gasoline |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CN201210189757.0A CN103483121B (en) | 2012-06-08 | 2012-06-08 | Method of preparing butadiene from gasoline |
Publications (2)
| Publication Number | Publication Date |
|---|---|
| CN103483121A CN103483121A (en) | 2014-01-01 |
| CN103483121B true CN103483121B (en) | 2015-02-11 |
Family
ID=49823784
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CN201210189757.0A Active CN103483121B (en) | 2012-06-08 | 2012-06-08 | Method of preparing butadiene from gasoline |
Country Status (1)
| Country | Link |
|---|---|
| CN (1) | CN103483121B (en) |
Citations (2)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| GB596879A (en) * | 1942-06-29 | 1948-01-13 | Koppers Co Inc | A method of pyrolyzing a hydro-aromatic compound |
| CN1549851A (en) * | 2001-07-18 | 2004-11-24 | ���з����о���˾ | Hydrogenation and dehydrogenation processes and catalysts therefor |
-
2012
- 2012-06-08 CN CN201210189757.0A patent/CN103483121B/en active Active
Patent Citations (2)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| GB596879A (en) * | 1942-06-29 | 1948-01-13 | Koppers Co Inc | A method of pyrolyzing a hydro-aromatic compound |
| CN1549851A (en) * | 2001-07-18 | 2004-11-24 | ���з����о���˾ | Hydrogenation and dehydrogenation processes and catalysts therefor |
Non-Patent Citations (2)
| Title |
|---|
| 杨利斌等,.延迟焦化汽油蒸汽热裂解性能实验室评价.《石化技术与应用》.2007,第25卷(第4期),第324-326页. * |
| 郝昭,."裂解汽油加氢装置的工业应用及优化".《中国优秀硕士学位论文全文数据库(电子期刊)工程科技I辑》.2009,(第04期),B019-40. * |
Also Published As
| Publication number | Publication date |
|---|---|
| CN103483121A (en) | 2014-01-01 |
Similar Documents
| Publication | Publication Date | Title |
|---|---|---|
| CN101935265B (en) | Liquefied gas catalytic pyrolysis process | |
| CN102443430B (en) | Method for preparing ethylene cracking material by hydrogenising liquefied petroleum gas | |
| CN101450884B (en) | High unsaturated hydrocarbon selective hydrogenation method in cracking gas | |
| CN101613621B (en) | Method for selective hydrogenation and separation for pyrolysis gas | |
| CN101451077B (en) | Selective hydrogenation method of acetylene hydrocarbon and dialkene in cracking gas | |
| CN103333040A (en) | Low energy consumption propylene production technology | |
| CN101423452B (en) | Selective hydrogenation method for fore-fraction high unsaturated hydrocarbons mixed phase | |
| CN103483121B (en) | Method of preparing butadiene from gasoline | |
| CN103483127B (en) | The method of divinyl is prepared by Aromatic raffinate | |
| CN104250192B (en) | The preparation method of a kind of propylene and divinyl | |
| CN103483126B (en) | Preparation method of butadiene | |
| JP6486367B2 (en) | Process for producing hydrocarbon products | |
| CN110041157B (en) | Method for improving yield of propylene prepared from methanol and prolonging service life of catalyst | |
| CN104250185B (en) | A kind of preparation method of low-carbon alkene | |
| CN1194073C (en) | Process for catalytic thermocracking of bydrocarbon by descending-bed reactor | |
| CN103467233B (en) | Method for preparing butadiene from gasoline | |
| CN103483122B (en) | Method of preparing butadiene from gasoline | |
| CN103483125B (en) | A kind of method being prepared divinyl by chain hydrocarbon | |
| CN109665933B (en) | Carbon four full-hydrogenation device and full-hydrogenation method | |
| CN103483123B (en) | Method of preparing butadiene from normal alkane | |
| CN103483128B (en) | Method of preparing butadiene from normal alkane | |
| CN104611002A (en) | Liquid-state petroleum hydrocarbon cracking method with characteristic of propylene yield increase and aromatic hydrocarbon yield increase | |
| CN103483124B (en) | A kind of preparation method of divinyl | |
| CN103483129B (en) | A kind of method being prepared divinyl by alkyl cyclopentane | |
| CN101870631A (en) | Method for producing low-carbon olefins by using petroleum hydrocarbon |
Legal Events
| Date | Code | Title | Description |
|---|---|---|---|
| C06 | Publication | ||
| PB01 | Publication | ||
| C10 | Entry into substantive examination | ||
| SE01 | Entry into force of request for substantive examination | ||
| C14 | Grant of patent or utility model | ||
| GR01 | Patent grant |