CN1884236A - Method for producing aromatic hydrocarbon - Google Patents
Method for producing aromatic hydrocarbon Download PDFInfo
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- CN1884236A CN1884236A CN 200510027027 CN200510027027A CN1884236A CN 1884236 A CN1884236 A CN 1884236A CN 200510027027 CN200510027027 CN 200510027027 CN 200510027027 A CN200510027027 A CN 200510027027A CN 1884236 A CN1884236 A CN 1884236A
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Abstract
The invention discloses a manufacturing method of aromatic hydrocarbons, which is characterized by the following: adopting twice gas-liquid separation; transmitting reacted product and materials in the first gas-liquid separating tank at higher temperature; cooling gas; transmitting the gas in the second gas-liquid separating tank; saving energy obviously; fitting for industrial manufacturing of aromatic hydrocarbons.
Description
Technical field
The present invention relates to be used for the method that aromatic hydrocarbons is produced.Specifically relate to toluene disproportionation and toluene and C
9And above aromatic hydrocarbons alkyl transfering process or selective disproportionation of toluene technology or benzene and carbon nine and above aromatic hydrocarbons alkyl transfering process thereof or carbon nine and above aromatic hydrocarbons thereof take off alkyl meth or Xylene isomerization process.
Background technology
P-Xylol is one of main basic organic of petrochemical industry, in numerous chemical production field such as chemical fibre, synthetic resins, agricultural chemicals, medicine, plastics purposes is widely arranged.Typical p-Xylol production method is that the dimethylbenzene that contains ethylbenzene of the thermodynamic(al)equilibrium that generates from the petroleum naphtha catalytic reforming is to separate or molecular sieve simulated moving bed adsorption separation (abbreviation fractionation by adsorption) technology by multistage cryogenic crystallization the C8 aronmatic, and p-Xylol is separated from the close with it isomer mixture of boiling point.And, often take C8 aronmatic (C for the processing of the dimethylbenzene of an ortho position and a position
8A) isomerization (abbreviation isomerization) technology makes it isomery and turns to p-Xylol.In order to increase production p-Xylol, utilize toluene disproportionation or toluene and carbon nine and above aromatic hydrocarbons (C thereof
9 +A) disproportionation and transalkylation (being called for short toluene disproportionation and transalkylation) reaction generates benzene and C
8A, thereby volume increase C
8A is the operational path that effectively increases production p-Xylol.In this operational path, toluene disproportionation wherein and transalkylation reaction and isomerization reaction all are the reactions of carrying out under hydro condition.
Up to now, in the world more typically, also the technology relevant with toluene disproportionation of comparative maturity has the traditional toluene disproportionation process of industrialized Tatoray of sixties Mo and realizes industrialized MSTDP selective disproportionation of toluene technology the late nineteen eighties.Selective disproportionation of toluene is a new way of producing p-Xylol, and toluene carries out the C that selective disproportionation generates benzene and high p-Xylol concentration
8A.
The document of p-Xylol production technique has USP2795629, USP4341914 etc.The technical process that representative USP4341914 proposes is for to carry out the aromatic hydrocarbons fractionation with reformate, toluene that obtains and C9 aromatic (C
9A) enter the Tatoray technique unit and carry out disproportionation and transalkylation reaction, resultant of reaction is toluene and C after separating
9A and part carbon+aromatic hydrocarbons (C
10A) circulation, benzene is as the product extraction, the C of generation
8A and the C that comes from isomerization unit
8A enters the p-Xylol tripping device together and isolates highly purified p-Xylol product, other C
8The A isomer carries out the xylol that xylene isomerization reaction obtains being in thermodynamic(al)equilibrium again to isomerization unit.
More than produce in the process integration of p-Xylol, no matter be toluene disproportionation unit or xylene isomerization unit, it all is the aromatic hydrocarbons conversion reaction of under the gaseous state hydro condition, carrying out, comprise aromatic hydrocarbons and hydrogen and lighter hydrocarbons etc. in the reaction product, reaction product at first with the reaction raw materials heat exchange after, temperature drops to about 80～200 ℃, under this temperature, also contain a large amount of aromatic hydrocarbons in the gaseous state, in order to reclaim aromatic hydrocarbons as far as possible, also must and be cooled to lower temperature through air cooler or the further condensation of watercooler, enter vapor-liquid separation tank then and carry out gas-liquid separation, separate the liquid aromatic hydrocarbons obtain and enter follow-up aromatic hydrocarbons fractionation unit, loop back reactive system after the isolated compressed machine supercharging of hydrogen that contains light hydrocarbon gas.Generally speaking, reaction product must satisfy the condition that can make most carbon six and above aromatic hydrocarbons thereof be condensed into liquid through air cooler or the cooled temperature of water cooler condensation, avoids aromatic hydrocarbons to lose in gas phase, and this temperature is 35～45 ℃ generally speaking.So, have only 35～45 ℃ from the isolated arene liquid of vapor-liquid separation tank, because temperature is too low, this logistics must be heated to certain temperature before entering separation column, heated temperatures is generally 100～200 ℃, this and through with the reaction raw materials heat exchange after the temperature of reaction product suitable.This shows, originally the higher reaction product with after the reaction raw materials heat exchange of temperature has been cooled to earlier and has carried out vapor-liquid separation about 35～45 ℃ again, obtain after the separation have only liquid about 35～45 ℃ to heat again to get back to quite high temperature after just can enter separation column, therefore, it is unreasonable that this process exists Btu utilization, the energy consumption problem of higher, the technical process aspect exists not enough.
Summary of the invention
Technical problem to be solved by this invention be in the conventional art when producing p-Xylol, the aromatic hydrocarbons conversion unit comprises toluene disproportionation and transalkylation, isomerization unit, methylbenzene shape selective disproportionation unit and benzene and C
9 +The A transalkylation, exist because flow process is unreasonable, reaction product with comparatively high temps is cooled to lower temperature earlier and carries out vapor-liquid separation, the liquid that separation obtains heats again gets back to higher temperature, cause problems such as energy consumption height, heat-exchange equipment load is big, a kind of new method that aromatic hydrocarbons is produced that is used for is provided.This method has can effectively reduce reaction product condensation cooling heat load, reduces follow-up separation column reboiler thermal load simultaneously, thus the characteristics that cut down the consumption of energy.
In order to solve the problems of the technologies described above, the technical solution used in the present invention is as follows: a kind of method that is used for aromatic hydrocarbons production may further comprise the steps:
A) temperature with 70～250 ℃ enters first vapor-liquid separation tank through hydrogen after the aromatic hydrocarbons conversion reaction and after being selected from mixture at least a in benzene, toluene, C8 aronmatic or carbon nine and the above aromatic hydrocarbons thereof and the heat exchange of aromatic hydrocarbons conversion reaction raw material, separate obtaining liquid stream and gas stream, gas stream enters b) step;
B) the isolated gas stream of first vapor-liquid separation tank enters the water cooler condensation and is cooled to 30～60 ℃ temperature, enter second vapor-liquid separation tank then, be circulated back to reactive system after the compressed machine supercharging of isolated gas, separate the liquid obtain and enter the aromatic hydrocarbons fractionating system after the liquid stream that obtains converges with separating from first vapor-liquid separation tank.
In technique scheme, the aromatic hydrocarbons conversion reaction comprises: toluene disproportionation and transalkylation reaction, toluene selective disproportionation reaction, xylene isomerization reaction, benzene and C9 aromatic transalkylation reaction and C9 aromatic dealkylation or the like.The operational condition of these technique unit reaction zones is as follows: pressure is 0.4～4Mpa, and temperature of reaction is 200～600 ℃, and hydrogen hydrocarbon mol ratio is 0.5～10, and the liquid weight air speed is 0.8～10 hour
-1The operational condition of first vapor-liquid separation tank is as follows: pressure is 0.1～5MPa, and temperature is 70～250 ℃; The operational condition of second gas separate pot is as follows: pressure is 0.1～5MPa, and temperature is 30～60 ℃.
Among the present invention, behind gaseous reaction products and reaction raw materials heat exchange after the aromatic hydrocarbons conversion, do not carry out the cooling of condensation further, carry out vapor-liquid separation but directly enter first vapor-liquid separation tank, only the higher gas of temperature that separation is obtained carries out the cooling of condensation further, enter second vapor-liquid separation tank then and carry out vapor-liquid separation, to reclaim the aromatic hydrocarbons in the gas as far as possible.Through after the above-mentioned flow process, isolated the higher liquid reacting product of most of temperature from first vapor-liquid separation tank, and bottom second separating tank, isolated the lower liquid reacting product of small portion temperature, these two strands of liquid are converged, the temperature of the product liquid that obtains is than higher, be generally about 70～200 ℃, so only need less heat just can be heated to it the operational requirement of knockout tower.As can be seen, such operational path had both reduced the condensation cooling heat of reaction product and had loaded, reduced the required thermal load of reacting by heating product again, reduced energy consumption and facility investment expense effectively, the ratio that energy consumption reduces is about 20～40%, has obtained better technical effect.
Description of drawings
The toluene disproportionation device technique flow process that Fig. 1 announces for USP2795629.
The toluene disproportionation device technique flow process that Fig. 2 announces for USP4341914.
Fig. 3 is used for the technical process of aromatic hydrocarbons conversion system for the present invention.
I is the feed for disproportionation pump among Fig. 1, Fig. 2 or Fig. 3, and II is a disproportionation reactor, and III is that first vapor-liquid separation tank, IV are second vapor-liquid separation tank, V is the 3rd vapor-liquid separation tank, and VI is a stripping tower, and VII is a toluene tower, VIII is a benzenol hydrorefining, and IX is a compressor, and X is the raw material surge tank, XI is a process furnace, XII is the reaction product air cooler, and XIII is the reaction product water cooler, and XIV is the stripping tower feed preheater, XV is a stripper reboiler, and XVI is reaction feed/effluent exchanger.1 is the disproportionation raw material, 2 is hydrogen make-up, 3 is circulating hydrogen, and 5 are reactor feed, and 6 is the reactor discharging, 7 are the reaction product behind process reaction feed/effluent exchanger, 8 is reaction product air cooler outlet material, and 9 is the outlet material of reaction product watercooler, and 10 are discharging gas, 11 is circulating hydrogen, 12 is the isolated liquid in vapour liquid separator IV bottom, and 13 is the isolated liquid in vapour liquid separator V bottom, and 14 is the isolated liquid in vapour liquid separator III bottom, 15 is the stripping tower charging, 16 is stripping tower still extraction liquid, and 17 is through cooled stripping tower extraction liquid, and 18 is the extraction of stripping cat head gas phase, 19 is the extraction of stripping cat head liquid phase, 20 is toluene, and 21 is toluene tower still extraction liquid, and 22 is dimethylbenzene, 23 is benzenol hydrorefining still extraction liquid, and 24 is the inlet logistics of stripping tower feed preheater in Fig. 3 flow process.
In the toluene disproportionation process flow process shown in Figure 1, disproportionation reaction product 6 is (not shown) after the condensation cooling, enter vapour liquid separator III, the isolated gas in top enters vapour liquid separator IV, the gas that rises contacts with the liquid 13 that descends from the top, the hydrocarbons that does not reclaim in the gas is absorbed, and is recycle hydrogen from container IV top expellant gas 11, turns back to behind the compressed machine among the reactor I.And the liquid 12 that discharge the bottom enters vapour liquid separator V, and gas 10 is discharged at the top, and bottom liquid 13 turns back to vapour liquid separator IV top.The liquid of separating from vapour liquid separator III 14 enters stripping tower VI, obtain logistics 19 from cat head, tower still extraction liquid 16 enters toluene tower VII, the toluene 20 that cat head obtains, circulation enters disproportionation reactor, and tower still extraction liquid 21 enters benzenol hydrorefining, and cat head obtains dimethylbenzene 22, and tower bottoms 23 is discharged.
In the toluene disproportionation process flow process shown in Figure 2, after disproportionation reaction product 6 and the reaction feed heat exchange, enter reaction product air cooler XII and reaction product water cooler XIII condensation cooling, enter vapour liquid separator III then, isolated portion gas 10 dischargings in top, most of gas enters circulating hydrogen compressor with hydrogen make-up, and the recycle hydrogen 3 after supercharging converges with raw material.Enter stripping tower VI after liquid of from vapour liquid separator III, separating 14 and 16 heat exchange of stripping tower still extraction liquid, isolate gaseous stream 18 and liquid phase stream 19 from cat head.
Aromatic hydrocarbons conversion process flow process of the present invention as shown in Figure 3.With the improvements of technical process illustrated in figures 1 and 2 be with the reaction feed heat exchange after disproportionation reaction product 7 do not carry out the cooling of condensation further, carry out vapor-liquid separation but directly enter the first vapour liquid separator III, the isolated gas 8 in top enters reaction product air cooler XII and reaction product water cooler XIII condensation cooling, and then enter the second vapour liquid separator IV, isolated portion gas 10 dischargings in top, most of gas enters circulating hydrogen compressor with hydrogen make-up, and the recycle hydrogen 3 after supercharging converges with raw material.Enter stripping tower VI after logistics 24 after converging from the higher liquid 14 of the isolated temperature of vapour liquid separator III with from the lower liquid 12 of the isolated temperature of vapour liquid separator IV and 16 heat exchange of stripping tower still extraction liquid, isolate gaseous stream 18 and liquid phase stream 19 from cat head.
The invention will be further elaborated below by embodiment.
Embodiment
[embodiment 1]
With toluene disproportionation and alkyl transfering process is example, the energy consumption situation when investigation the present invention carries out toluene disproportionation and transalkylation reaction by the technical process shown in the accompanying drawing 3.
Toluene and C
9 +A raw material 1 is with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, enter reactor and carry out toluene disproportionation and transalkylation reaction, generate C8 aronmatic and benzene, enter first vapour liquid separator after reaction product 6 and the reaction raw materials heat exchange and carry out vapor-liquid separation, the isolated gas that contains than polyaromatic enters second vapor-liquid separation tank and carries out vapor-liquid separation once more after condensation is cooled to 38 ℃.Two isolated liquid reacting products of vapor-liquid separation tank converge forms logistics 24, enters stripping tower and separates then with after 16 heat exchange of stripping tower tower bottoms.
Table 1 is the flow and the service temperature of each logistics, and table 2 is thermal loads of each interchanger.
Table 1 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 24 | 15 |
| Flow, kilogram/hour | 197454 | 82749 | 125602 | 125602 |
| Temperature, ℃ | 88 | 38 | 85 | 172 |
Table 2 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 5.97 | 9.10 |
The result shows, utilization the technology of the present invention, and reaction product condensation cooling heat load is 5.97MW, and the total condensation cooling heat load in the comparative example 1 is to have descended 27.5% by 8.23MW.The temperature of vapor-liquid separation tank outlet liquid has been brought up to 85 ℃ from 38 ℃, and the feeding temperature of stripping tower has been brought up to 172 ℃ from 139 ℃, and also to be the stripper reboiler thermal load drop to 9.10MW from the 12.23MW of comparative example 1 for this, and 25.6% reason has descended.Thermal load descends and has not only reduced the general facilities consumption, has reduced production cost, has improved productivity effect, but also has reduced the facility investment expense.Therefore, it is unreasonable that the present invention has overcome the Btu utilization that exists in the conventional art, causes the energy consumption problem of higher, and a kind of new toluene disproportionation and transalkylation method are provided.
[embodiment 2]
Be example still with toluene disproportionation and alkyl transfering process, wherein the flow of reaction product stream 7 is identical with embodiment 1, but temperature is brought up to 126 ℃ from 88 ℃, the energy consumption situation when investigating the present invention and carry out toluene disproportionation and transalkylation reaction by the technical process shown in the accompanying drawing 3 with this.
Table 3 is flow and service temperatures of each logistics, and table 4 is thermal loads of each interchanger.
Table 3 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 24 | 15 |
| Flow, kilogram/hour | 197454 | 82749 | 125602 | 125602 |
| Temperature, ℃ | 126 | 38 | 102 | 182 |
Table 4 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 13.31 | 7.91 |
The result shows, utilization the technology of the present invention, and reaction product condensation cooling heat load is 13.31MW, and the total condensation cooling heat load in the comparative example 2 is to have descended 20.6% by 16.76MW.The temperature of vapor-liquid separation tank outlet liquid has been brought up to 102 ℃ from 38 ℃, and the feeding temperature of stripping tower has been brought up to 182 ℃ from 139 ℃, and also to be the stripper reboiler thermal load drop to 7.91MW from the 12.23MW of comparative example 2 for this, and 35.3% reason has descended.Thermal load descends and has not only reduced the general facilities consumption, has reduced production cost, has improved productivity effect, but also has reduced the facility investment expense.Therefore, it is unreasonable that the present invention has overcome the Btu utilization that exists in the conventional art, causes the energy consumption problem of higher, and a kind of new toluene disproportionation and transalkylation method are provided.
[embodiment 3]
Selecting the type disproportionation processes with toluene is example, investigates the present invention and carries out the energy consumption situation of toluene when selecting the type disproportionation reaction by the technical process shown in the accompanying drawing 3.
Methylbenzene raw material 1 is with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, entering reactor carries out toluene and selects the type disproportionation reaction, generation contains the C8 aronmatic and the benzene of high density p-Xylol, enter first vapour liquid separator after reaction product 6 and the reaction raw materials heat exchange and carry out vapor-liquid separation, the isolated gas that contains than polyaromatic enters second vapor-liquid separation tank and carries out vapor-liquid separation once more after condensation is cooled to 38 ℃.Two isolated liquid reacting products of vapor-liquid separation tank converge forms logistics 24, enters stripping tower and separates then with after 16 heat exchange of stripping tower tower bottoms.
Table 5 is flow and service temperatures of each logistics, and table 6 is thermal loads of each interchanger.
Table 5 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 24 | 15 |
| Flow, kilogram/hour | 144204 | 93484 | 123857 | 123857 |
| Temperature, ℃ | 146 | 38 | 87 | 183 |
Table 6 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 15.19 | 6.17 |
The result shows, utilization the technology of the present invention, and reaction product condensation cooling heat load is 15.19MW, and the total condensation cooling heat load in the comparative example 3 is to have descended 16.2% by 18.12MW.The temperature of vapor-liquid separation tank outlet liquid has been brought up to 87 ℃ from 38 ℃, and the feeding temperature of stripping tower has been brought up to 183 ℃ from 145 ℃, and also to be the stripper reboiler thermal load drop to 6.17MW from the 9.47MW of comparative example 3 for this, and 34.8% reason has descended.Thermal load descends and has not only reduced the general facilities consumption, has reduced production cost, has improved productivity effect, but also has reduced the facility investment expense.Therefore, the present invention overcome toluene in the past select exist in the type disproportionation processes cause the energy consumption problem of higher owing to technical process is unreasonable, it is reasonable to have a Btu utilization, the advantage that energy consumption is lower can be used for toluene and selects in the type disproportionation processes.
[embodiment 4]
Still selecting the type disproportionation processes with toluene is example, and wherein the flow of reaction product stream 7 is identical with embodiment 3, but temperature is reduced to 88 ℃ from 146 ℃, investigates the present invention with this and carries out the energy consumption situation of toluene when selecting the type disproportionation reaction by the technical process shown in the accompanying drawing 3.
Table 7 is flow and service temperatures of each logistics, and table 8 is thermal loads of each interchanger.
Table 7 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 24 | 15 |
| Flow, kilogram/hour | 144204 | 35882 | 123857 | 123857 |
| Temperature, ℃ | 89 | 38 | 85 | 181 |
Table 8 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 3.63 | 6.08 |
The result shows, utilization the technology of the present invention, and reaction product condensation cooling heat load is 3.63MW, and the total condensation cooling heat load in the comparative example 4 is to have descended 43.2% by 6.39MW.The temperature of vapor-liquid separation tank outlet liquid has been brought up to 85 ℃ from 38 ℃, and the feeding temperature of stripping tower has been brought up to 181 ℃ from 145 ℃, and also to be the stripper reboiler thermal load drop to 6.08MW from the 9.47MW of comparative example 4 for this, and 35.8% reason has descended.Thermal load descends and has not only reduced the general facilities consumption, has reduced production cost, has improved productivity effect, but also has reduced the facility investment expense.Therefore, toluene is selected type disproportionation device and is adopted technology of the present invention, can significantly cut down the consumption of energy, and it is reasonable to have a Btu utilization, the advantage that energy consumption is lower.
[embodiment 5]
With benzene and C9 aromatic alkyl transfering process is example, investigates the present invention and carries out C9 aromatic and the benzene energy consumption situation when carrying out transalkylation reaction by the technical process shown in the accompanying drawing 3.
Benzene and C
9 +A raw material 1 is with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, enter reactor and carry out transalkylation reaction, generate toluene and C8 aronmatic, enter first vapour liquid separator after reaction product 6 and the reaction raw materials heat exchange and carry out vapor-liquid separation, enter second vapor-liquid separation tank after the isolated condensation of gas that contains higher concentration aromatic hydrocarbons is cooled to 38 ℃ and carry out vapor-liquid separation once more.Two isolated liquid of vapor-liquid separation tank converge forms logistics 24, enters stripping tower and separates then with after 16 heat exchange of stripping tower tower bottoms.
Table 9 is flow and service temperatures of each logistics, and table 10 is thermal loads of each interchanger.
Table 9 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 24 | 15 |
| Flow, kilogram/hour | 237976 | 13337 | 139164 | 139164 |
| Temperature, ℃ | 102 | 38 | 88 | 167 |
Table 10 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 12.04 | 16.11 |
The result shows, utilization the technology of the present invention, and reaction product condensation cooling heat load is 12.04MW, and the total condensation cooling heat load in the comparative example 5 is to have descended 30.9% by 17.43MW.And the stripping tower thermal load is 16.11MW, compares with the 20.76MW of comparative example 5, has descended 22.4%.Therefore, the present invention can be applicable in benzene and the C9 aromatic alkyl transfering process, and it is reasonable to have a Btu utilization, the advantage that energy consumption is lower.
[comparative example 1]
Adopt the raw material identical, the energy consumption situation when toluene disproportionation and toluene and carbon nine and above aromatic hydrocarbons transalkylation reaction thereof are carried out in investigation technical process as shown in Figure 2 with embodiment 1.
Toluene and carbon nine and above aroamtic hydrocarbon raw material 1 thereof are with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, enter reactor and carry out toluene disproportionation and toluene and carbon nine and above aromatic hydrocarbons transalkylation reaction thereof, generate C8 aronmatic and benzene, reaction product 6 elder generations and reaction raw materials heat exchange continue condensation then and are cooled to 38 ℃, enter vapor-liquid separation tank and carry out vapor-liquid separation.Entering stripping tower after isolated liquid reacting product 14 and 16 heat exchange of stripping tower tower bottoms separates.
Table 11 is flow and service temperatures of each logistics, and the thermal load of each heat-exchange equipment is listed in the table 12.
Table 11 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 14 | 15 |
| Flow, kilogram/hour | 197454 | 197454 | 125602 | 125602 |
| Temperature, ℃ | 88 | 38 | 38 | 139 |
Table 12 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 8.23 | 12.23 |
As seen from Table 12, total condensation cooling heat load is 8.23MW, and total heat hot load is 12.23MW.
[comparative example 2]
Adopt the raw material identical, the energy consumption situation when toluene disproportionation and toluene and carbon nine and above aromatic hydrocarbons transalkylation reaction thereof are carried out in investigation technical process as shown in Figure 2 with embodiment 2.Table 13 is flow and service temperatures of each logistics, and the thermal load of each heat-exchange equipment is listed in the table 14.
Table 13 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 14 | 15 |
| Flow, kilogram/hour | 197454 | 197454 | 125602 | 125602 |
| Temperature, ℃ | 126 | 38 | 38 | 139 |
Table 14 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 16.76 | 12.23 |
As seen from Table 14, total cooling heat load is 16.76MW, and total heat hot load is 12.23MW.
[comparative example 3]
Adopt the raw material identical, the energy consumption situation when toluene selective disproportionation reaction is carried out in investigation technical process as shown in Figure 2 with embodiment 3.
Methylbenzene raw material 1 is with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, entering reactor carries out toluene and selects the type disproportionation reaction, generation contains the C8 aronmatic and the benzene of high density p-Xylol, reaction product 6 elder generations and reaction raw materials heat exchange continue condensation then and are cooled to 38 ℃, enter vapor-liquid separation tank and carry out vapor-liquid separation.Entering stripping tower after isolated liquid reacting product 14 and 16 heat exchange of stripping tower tower bottoms separates.
Table 15 is flow and service temperatures of each logistics, and the thermal load of each heat-exchange equipment is listed in the table 16.
Table 15 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 14 | 15 |
| Flow, kilogram/hour | 144204 | 144204 | 123857 | 123857 |
| Temperature, ℃ | 146 | 38 | 38 | 145 |
Table 16 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 18.12 | 9.47 |
As seen from Table 16, total cooling heat load is 18.12MW, and total heat hot load is 9.47MW.
[comparative example 4]
Adopt the raw material identical, the energy consumption situation when toluene selective disproportionation reaction is carried out in investigation technical process as shown in Figure 2 with embodiment 4.Table 17 is flow and service temperatures of each logistics, and the thermal load of each heat-exchange equipment is listed in the table 18.
Table 17 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 14 | 15 |
| Flow, kilogram/hour | 144204 | 144204 | 124148 | 124148 |
| Temperature, ℃ | 88 | 38 | 38 | 145 |
Table 18 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 6.39 | 9.47 |
As seen from Table 18, total cooling heat load is 6.39MW, and total heat hot load is 9.47MW.
[comparative example 5]
Adopt the raw material identical, the energy consumption situation when benzene and C9 aromatic transalkylation reaction are carried out in investigation technical process as shown in Figure 2 with embodiment 5.
Benzene and carbon nine and above aroamtic hydrocarbon raw material 1 thereof are with after circulating hydrogen 3 mixes, with reaction product 6 heat exchange, be heated to temperature of reaction through process furnace again, enter reactor and carry out carbon nine and above aromatic hydrocarbons and benzene transalkylation reaction, generate C8 aronmatic and toluene, reaction product 6 elder generations and reaction raw materials heat exchange continue condensation then and are cooled to 38 ℃, enter vapor-liquid separation tank and carry out vapor-liquid separation.Entering stripping tower after isolated liquid reacting product 14 and 16 heat exchange of stripping tower tower bottoms separates.
Table 19 is flow and service temperatures of each logistics, and the thermal load of each heat-exchange equipment is listed in the table 20.
Table 19 logistics flux and service temperature
| The logistics numbering | 7 | 9 | 14 | 15 |
| Flow, kilogram/hour | 237976 | 237967 | 139164 | 139164 |
| Temperature, ℃ | 102 | 38 | 38 | 132 |
Table 20 exchanger heat load meter
| The reaction product condensate cooler | Stripper reboiler | |
| Thermal load, MW | 17.43 | 20.76 |
As seen from Table 20, total cooling heat load is 17.43MW, and total heat hot load is 20.76MW.
Claims (3)
1, a kind of method that is used for aromatic hydrocarbons production may further comprise the steps:
A) temperature with 70～250 ℃ enters first vapor-liquid separation tank through hydrogen after the aromatic hydrocarbons conversion reaction and after being selected from mixture at least a in benzene, toluene, C8 aronmatic or carbon nine and the above aromatic hydrocarbons thereof and the heat exchange of aromatic hydrocarbons conversion reaction raw material, separate obtaining liquid stream and gas stream, gas stream enters b) step;
B) the isolated gas stream of first vapor-liquid separation tank enters the water cooler condensation and is cooled to 30～60 ℃ temperature, enter second vapor-liquid separation tank then, Returning reacting system after the compressed machine supercharging of isolated gas separates the liquid obtain and enters the aromatic hydrocarbons fractionating system after the liquid stream that obtains converges with separating from first vapor-liquid separation tank.
2, according to the described method that is used for aromatic hydrocarbons production of claim 1, the working pressure that it is characterized in that first vapor-liquid separation tank is 0.1～5MPa, and service temperature is 70～250 ℃.
3, according to the described method that is used for aromatic hydrocarbons production of claim 1, the working pressure that it is characterized in that second vapor-liquid separation tank is 0.1～5MPa, and service temperature is 30～60 ℃.
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| CNB2005100270270A CN100335449C (en) | 2005-06-22 | 2005-06-22 | Method for producing aromatic hydrocarbon |
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| CNB2005100270270A CN100335449C (en) | 2005-06-22 | 2005-06-22 | Method for producing aromatic hydrocarbon |
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| CN100335449C CN100335449C (en) | 2007-09-05 |
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Cited By (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN105712814A (en) * | 2014-12-05 | 2016-06-29 | 中国石油化工股份有限公司 | Improved separation method |
| CN105712815A (en) * | 2014-12-05 | 2016-06-29 | 中国石油化工股份有限公司 | Energy saving separation method |
| CN110325494A (en) * | 2016-12-28 | 2019-10-11 | 环球油品公司 | Heat in alkylation process is reduced |
| CN114100558A (en) * | 2021-11-26 | 2022-03-01 | 东营威联化学有限公司 | Equipment for producing benzene and xylene by toluene disproportionation |
Family Cites Families (3)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US4041091A (en) * | 1976-02-19 | 1977-08-09 | Uop Inc. | Fractionation of aromatic streams |
| US4341914A (en) * | 1980-12-22 | 1982-07-27 | Uop Inc. | Transalkylation process with recycle of C10 hydrocarbons |
| CN1164541C (en) * | 2001-10-22 | 2004-09-01 | 中国石油化工股份有限公司 | Toluene selective disproportion and toluene and C9 and its above aromatic disproportion and alkyl transfusion method |
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Cited By (5)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| CN105712814A (en) * | 2014-12-05 | 2016-06-29 | 中国石油化工股份有限公司 | Improved separation method |
| CN105712815A (en) * | 2014-12-05 | 2016-06-29 | 中国石油化工股份有限公司 | Energy saving separation method |
| CN110325494A (en) * | 2016-12-28 | 2019-10-11 | 环球油品公司 | Heat in alkylation process is reduced |
| CN110325494B (en) * | 2016-12-28 | 2022-08-16 | 环球油品公司 | Heat reduction in alkylation processes |
| CN114100558A (en) * | 2021-11-26 | 2022-03-01 | 东营威联化学有限公司 | Equipment for producing benzene and xylene by toluene disproportionation |
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| Publication number | Publication date |
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| CN100335449C (en) | 2007-09-05 |
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