CN1157280A - Light hydrocarbon separation method capable of raising ethylene recovery - Google Patents
Light hydrocarbon separation method capable of raising ethylene recovery Download PDFInfo
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- CN1157280A CN1157280A CN96120253A CN96120253A CN1157280A CN 1157280 A CN1157280 A CN 1157280A CN 96120253 A CN96120253 A CN 96120253A CN 96120253 A CN96120253 A CN 96120253A CN 1157280 A CN1157280 A CN 1157280A
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- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 title claims abstract description 82
- 239000005977 Ethylene Substances 0.000 title claims abstract description 52
- 229930195733 hydrocarbon Natural products 0.000 title claims abstract description 20
- 150000002430 hydrocarbons Chemical class 0.000 title claims abstract description 20
- 238000011084 recovery Methods 0.000 title claims abstract description 15
- 238000000926 separation method Methods 0.000 title claims abstract description 15
- 239000004215 Carbon black (E152) Substances 0.000 title claims abstract description 8
- 238000000034 method Methods 0.000 claims abstract description 116
- 230000008569 process Effects 0.000 claims abstract description 109
- VNWKTOKETHGBQD-UHFFFAOYSA-N methane Chemical compound C VNWKTOKETHGBQD-UHFFFAOYSA-N 0.000 claims abstract description 107
- 239000007788 liquid Substances 0.000 claims abstract description 32
- 238000010521 absorption reaction Methods 0.000 claims abstract description 21
- 238000001816 cooling Methods 0.000 claims abstract description 19
- 238000007701 flash-distillation Methods 0.000 claims abstract description 8
- 238000005520 cutting process Methods 0.000 claims abstract description 4
- 238000007600 charging Methods 0.000 claims description 30
- 239000007791 liquid phase Substances 0.000 claims description 29
- OKTJSMMVPCPJKN-UHFFFAOYSA-N Carbon Chemical compound [C] OKTJSMMVPCPJKN-UHFFFAOYSA-N 0.000 claims description 19
- 229910052799 carbon Inorganic materials 0.000 claims description 19
- HSFWRNGVRCDJHI-UHFFFAOYSA-N alpha-acetylene Natural products C#C HSFWRNGVRCDJHI-UHFFFAOYSA-N 0.000 claims description 14
- 125000002534 ethynyl group Chemical group [H]C#C* 0.000 claims description 12
- 239000003795 chemical substances by application Substances 0.000 claims description 10
- 241000282326 Felis catus Species 0.000 claims description 9
- 239000006096 absorbing agent Substances 0.000 claims description 9
- 238000005984 hydrogenation reaction Methods 0.000 claims description 9
- OTMSDBZUPAUEDD-UHFFFAOYSA-N Ethane Chemical compound CC OTMSDBZUPAUEDD-UHFFFAOYSA-N 0.000 claims description 8
- 238000007599 discharging Methods 0.000 claims description 6
- 238000004134 energy conservation Methods 0.000 claims description 4
- 238000011144 upstream manufacturing Methods 0.000 claims description 2
- 239000000463 material Substances 0.000 abstract description 3
- 230000006835 compression Effects 0.000 abstract 1
- 238000007906 compression Methods 0.000 abstract 1
- 239000007789 gas Substances 0.000 description 34
- 239000000047 product Substances 0.000 description 19
- 238000005265 energy consumption Methods 0.000 description 13
- 239000012071 phase Substances 0.000 description 11
- 125000004805 propylene group Chemical group [H]C([H])([H])C([H])([*:1])C([H])([H])[*:2] 0.000 description 10
- QQONPFPTGQHPMA-UHFFFAOYSA-N propylene Natural products CC=C QQONPFPTGQHPMA-UHFFFAOYSA-N 0.000 description 9
- 239000001257 hydrogen Substances 0.000 description 8
- 229910052739 hydrogen Inorganic materials 0.000 description 8
- 238000010992 reflux Methods 0.000 description 7
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 description 6
- 239000004234 Yellow 2G Substances 0.000 description 6
- 239000003507 refrigerant Substances 0.000 description 6
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 description 6
- 239000000470 constituent Substances 0.000 description 5
- 238000005261 decarburization Methods 0.000 description 5
- 230000000694 effects Effects 0.000 description 4
- 239000000203 mixture Substances 0.000 description 4
- 239000002151 riboflavin Substances 0.000 description 4
- 239000004229 Alkannin Substances 0.000 description 3
- ATUOYWHBWRKTHZ-UHFFFAOYSA-N Propane Chemical compound CCC ATUOYWHBWRKTHZ-UHFFFAOYSA-N 0.000 description 3
- 239000004176 azorubin Substances 0.000 description 3
- 230000008901 benefit Effects 0.000 description 3
- 239000001679 citrus red 2 Substances 0.000 description 3
- 239000000284 extract Substances 0.000 description 3
- 238000012545 processing Methods 0.000 description 3
- 230000000630 rising effect Effects 0.000 description 3
- 239000004149 tartrazine Substances 0.000 description 3
- 238000012546 transfer Methods 0.000 description 3
- KAKZBPTYRLMSJV-UHFFFAOYSA-N Butadiene Chemical group C=CC=C KAKZBPTYRLMSJV-UHFFFAOYSA-N 0.000 description 2
- 239000004235 Orange GGN Substances 0.000 description 2
- 239000004231 Riboflavin-5-Sodium Phosphate Substances 0.000 description 2
- 230000008859 change Effects 0.000 description 2
- 238000009833 condensation Methods 0.000 description 2
- 230000005494 condensation Effects 0.000 description 2
- 230000006837 decompression Effects 0.000 description 2
- 238000013461 design Methods 0.000 description 2
- 150000002431 hydrogen Chemical class 0.000 description 2
- 230000010354 integration Effects 0.000 description 2
- 238000011027 product recovery Methods 0.000 description 2
- 230000009467 reduction Effects 0.000 description 2
- 238000001179 sorption measurement Methods 0.000 description 2
- 239000004230 Fast Yellow AB Substances 0.000 description 1
- 230000009471 action Effects 0.000 description 1
- FFBHFFJDDLITSX-UHFFFAOYSA-N benzyl N-[2-hydroxy-4-(3-oxomorpholin-4-yl)phenyl]carbamate Chemical compound OC1=C(NC(=O)OCC2=CC=CC=C2)C=CC(=C1)N1CCOCC1=O FFBHFFJDDLITSX-UHFFFAOYSA-N 0.000 description 1
- 238000004364 calculation method Methods 0.000 description 1
- 239000004106 carminic acid Substances 0.000 description 1
- 238000003889 chemical engineering Methods 0.000 description 1
- 239000002826 coolant Substances 0.000 description 1
- 238000005336 cracking Methods 0.000 description 1
- 230000007423 decrease Effects 0.000 description 1
- 230000018044 dehydration Effects 0.000 description 1
- 238000006297 dehydration reaction Methods 0.000 description 1
- 238000005516 engineering process Methods 0.000 description 1
- 125000000816 ethylene group Chemical group [H]C([H])([*:1])C([H])([H])[*:2] 0.000 description 1
- 238000000605 extraction Methods 0.000 description 1
- 238000005194 fractionation Methods 0.000 description 1
- 239000008246 gaseous mixture Substances 0.000 description 1
- 230000006872 improvement Effects 0.000 description 1
- 238000004519 manufacturing process Methods 0.000 description 1
- 230000004048 modification Effects 0.000 description 1
- 238000012986 modification Methods 0.000 description 1
- 239000001294 propane Substances 0.000 description 1
- 239000004173 sunset yellow FCF Substances 0.000 description 1
- 239000013589 supplement Substances 0.000 description 1
- 238000009834 vaporization Methods 0.000 description 1
- 230000008016 vaporization Effects 0.000 description 1
- 238000005303 weighing Methods 0.000 description 1
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0204—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the feed stream
- F25J3/0219—Refinery gas, cracking gas, coke oven gas, gaseous mixtures containing aliphatic unsaturated CnHm or gaseous mixtures of undefined nature
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- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0233—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of CnHm with 1 carbon atom or more
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0238—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of CnHm with 2 carbon atoms or more
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0242—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of CnHm with 3 carbon atoms or more
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- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J3/00—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification
- F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
- F25J3/0228—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream
- F25J3/0252—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream characterised by the separated product stream separation of hydrogen
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- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2200/00—Processes or apparatus using separation by rectification
- F25J2200/74—Refluxing the column with at least a part of the partially condensed overhead gas
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- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2200/00—Processes or apparatus using separation by rectification
- F25J2200/76—Refluxing the column with condensed overhead gas being cycled in a quasi-closed loop refrigeration cycle
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2205/00—Processes or apparatus using other separation and/or other processing means
- F25J2205/02—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum
- F25J2205/04—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum in the feed line, i.e. upstream of the fractionation step
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2205/00—Processes or apparatus using other separation and/or other processing means
- F25J2205/30—Processes or apparatus using other separation and/or other processing means using a washing, e.g. "scrubbing" or bubble column for purification purposes
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- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2210/00—Processes characterised by the type or other details of the feed stream
- F25J2210/12—Refinery or petrochemical off-gas
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- F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
- F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
- F25J—LIQUEFACTION, SOLIDIFICATION OR SEPARATION OF GASES OR GASEOUS OR LIQUEFIED GASEOUS MIXTURES BY PRESSURE AND COLD TREATMENT OR BY BRINGING THEM INTO THE SUPERCRITICAL STATE
- F25J2215/00—Processes characterised by the type or other details of the product stream
- F25J2215/62—Ethane or ethylene
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Abstract
The light hydrocarbon separation method for raising ethylene recovery includes the following steps: the gas and liquid obtained after compression, cooling and flash distillation oflight hydrocarbon are respectively fed into high-pressure deethanizer to make non-clearity cutting of C2-fraction; the bottom product of the high-pressure deethanizer is fed into low-pressure deethanizer to implement the separation of C2-fraction and C3-fraction, and the top product of the high-pressure deethanizer is passed through the processes of step-by-step cooling and flash distillation so as to obtain the liquid as feeding material of demethanizer, and the gas discharged from outlet of gas-liquid separating tank is cooled, and fed itno methane absorption tower, and 99.5% ethylene in the gas can be absorbed.
Description
The invention belongs to field of chemical engineering, relate in particular to the improvement of the recovery ethene method of ethylene unit and other lighter hydrocarbons processing unit (plant)s.
Well-known lighter hydrocarbons are meant the hydrocarbon cracking product that has removed carbon five or carbon four above heavy constituent, and its composition that also can refer to other sources mainly comprises the mixture of carbon five following hydrocarbon such as hydrogen, methane, ethene, ethane, propylene, propane.Lighter hydrocarbons need further processing, are separated into products such as ethene, propylene.Lighter hydrocarbons separate general employing rectifying and flash distillation.The index of weighing the separation method quality is energy consumption, investment and product recovery rate, the rate of recovery of major product ethene especially, and these three indexs are conflicting often: as require the product recovery rate height then usually to need higher energy consumption and/or investment.An outstanding separation method is to take into account the requirement of three aspects simultaneously, thereby makes production cost minimum.
The lighter hydrocarbons sepn process is broadly divided into demethanizing system, decarburization two system and decarburization three system's three parts, the loss of ethene mainly occurs in demethanizing system and decarburization two system, and the ethylene loss that occurs in the demethanizing system accounts for more than 60% of ethene total losses usually, two of the energy consumption maximum towers are ethylene rectification tower and demethanizing tower in the sepn process, therefore will fully pay attention to the design of demethanizing system.Multiple separation process is arranged at present, and wherein a large amount of a kind of demethanizing system flows (to call flow process one in the following text) that adopt are referring to figure one.This flow process is lighter hydrocarbons process compressor C101, and water cooler E101 enters knockout drum F101 behind moisture eliminator Y101 and the water cooler E102.The liquid of F101 outlet enters demethanizing tower D101, the gas of outlet passes through what cooling and flash distillation again, the outlet liquid of knockout drum F102, F103 and F104 still enters demethanizing tower, the gas of F104 outlet is the ethene of methane, hydrogen and the constraint that balanced each other, this gas continues by interchanger E109, be cooled to and enter last step knockout drum F105 about-165 ℃, exit gas mainly is a hydrogen, and outlet liquid mainly is methane.The required cold of E109 is provided by hydrogen and post-decompression methane.The low-pressure methane and the hydrogen that go out E109 enter preceding what interchanger E107～E103, also enter E102 sometimes, and continuing provides cold as low-temperature receiver.In order to reach the temperature out of F104, generally between-127 ℃～-130 ℃, except by the gaseous products methane of demethanizer column overhead after decompression also as low-temperature receiver, still need and extract part methane product low-temperature receiver as a supplement out, control the temperature of F104 by the quantity of extracting liquid phase methane out with liquid phase state.The separation requirement of demethanizing tower is that bottom product removes methane, overhead product removes ethene to set quota, and the method for control is with the conventional rectification tower.Demethanizing tower bottom product is carbon two and above cut, delivers to deethanizing column D104.The overhead product C-2-fraction of D104 enters ethylene rectification tower after removing acetylene through acetylene hydrogenation reactor R101, obtains major product ethene from the top of this tower, and bottom product is an ethane.The deethanizing column bottoms is carbon three or more last running, enters decarburization three systems.
By above narration as can be known, the ethylene loss of demethanizing system occurs in three places, i.e. the cat head gas phase discharging C of gas phase discharging B, the demethanizing tower D101 of A, F104, the cat head liquid phase discharging of demethanizing tower D101
A is ordered is lost in the temperature that depends on F104 under certain feed composition, the lower then ethylene loss of temperature is fewer, but the condenser duty of demethanizing tower is bigger, and the restriction that the temperature of this point also is subjected to balancing each other and heat transfer temperature difference requires, and temperature generally is not less than-130 ℃.
The ethylene loss that B, C are 2 depends on the reflux ratio of demethanizing tower and the extraction amount of liquid phase methane, and reduce this ethylene loss of 2 must be cost to increase energy consumption, and this point is self-evident.
Patent of invention 90101957.7 has been narrated a kind of technology that reclaims ethene from contain gaseous mixture such as methane, ethane and ethene.Equipment that a series of heat transfers of this process using and mass transfer carry out simultaneously and plural demethanizing tower, this method have very high ethylene recovery rate, but energy consumption and facility investment are not necessarily low.
Patent of invention ZL92100471.0 is that the energy consumption of double tower front-end deethanization flow process is lower, but the ethylene recovery rate is lower than patent of invention 90101957.7.
The objective of the invention is to improve the load of ethylene recovery rate, reduction demethanizing tower and cut down the consumption of energy.
The present invention is a kind of energy-conservation, ethylene recovery rate height and the little light hydrocarbon separating method of demethanizing tower load, flow process is charging after overdraft, cooling off and remove moisture content, further cooling and flash distillation obtain gas and liquid and enter the high pressure deethanizing column respectively, carry out the non-clear cutting of carbon two fractions.High pressure deethanizing column bottoms goes the low pressure deethanizing column to carry out separating of carbon two and c 3 fraction.The overhead product of high pressure deethanizing column obtains liquid more than one as the charging of demethanizing tower through cooling and flash distillation step by step, gas further enters an absorption tower after the cooling, remove to absorb ethene in the inlet gas with liquid phase methane, absorption tower outlet liquid returns demethanizing tower as the top charging.The demethanizing tower bottoms is carbon containing three fractions not, remove acetylene hydrogenation reactor or ethylene rectification tower.Ethylene rectification tower is removed in the top discharge of low pressure deethanizing column after removing acetylene, as second burst of charging of this tower, still liquid goes to decarburization three systems.The feature of flow process is to have the cat head discharging only to contain the high pressure deethanizing column of part carbon two fractions in the upstream of demethanizing tower and is the methane absorber of absorption agent with methane.The temperature of methane adsorption agent is-120 ℃～-145 ℃ scopes, and suitable temperature is-130 ℃--and 140 ℃.
When demethanizing tower was under high pressure operated, figure two was seen in technical process of the present invention.Lighter hydrocarbons are through compressor C101, dryer feed water cooler E101.Moisture eliminator V101 and moisture eliminator aftercooler are cooled to 2 ℃--and 25 ℃, air inlet liquid separating tank is separated into gas phase and liquid phase.Knockout drum can be an one-level, and what F101 as shown in the figure also can carry out and separate in process of cooling, obtain one above liquid phase.The paramount pressure-off ethane tower of liquid phase D102, gas phase enters D102 after further being cooled to-30 ℃～-37 ℃, carry out the non-clear cutting of C-2-fraction at D102, overhead product is the C-2-fraction of whole methane, hydrogen and 30%-70% in the charging, and bottom product is remaining C-2-fraction and more heavy constituent.The D102 overhead product is through boil among the D101 device E112 and interchanger E103～E106 cooling and partial condensation, obtains two strands of liquid in F102 and F103 flash distillation and enters demethanizing tower as charging.The gas of F103 outlet is cooled between-105 ℃～-121 ℃ through E107 and enters absorption tower D103, be used for from the demethanizing tower, temperature be liquid phase methane between-130 ℃～-140 ℃ as absorption agent, also the liquid phase methane in available other sources absorbs the ethene in the F103 exit gas get off as absorption agent.Specific absorption can be reached about 99.5% by gas-liquid phase temperature that enters the absorption tower and the decision of liquid phase quantity.Enter the top of demethanizing tower as first burst of charging of demethanizing tower from the absorption agent of D103 outlet at bottom.Still liquid carbon containing three cuts not at the bottom of demethanizing tower (D101) tower enter acetylene hydrogenation reactor (R101B) and ethylene rectification tower D105 and obtain ethene and ethane.
When demethanizing tower in low pressure, basic identical shown in technical process of the present invention and the figure two when for example operating in the 0.5-1.5MPA scope, difference mainly is the integrated mode difference of heat.Owing to removed part C-2-fraction and all heavy constituent of carbon more than three at the high pressure deethanizing column, the load of demethanizing tower stripping section is descended, required rising quantity of steam reduces.Utilize this condition, can optimize the hot state (q value) of demethanizing tower charging, promptly reduce the liquefied fraction of charging, adopt the liquid phase feeding of demethanizing tower is gone to cool off the method for demethanizing tower charging after vacuum flashing, make originally the cold of-62 ℃ of left and right sides grades that reclaim at the demethanizing tower reboiler be increased to the colds of grade for-101 ℃ and-135 ℃, thereby reduced energy consumption, figure six has represented a kind of so hot integration mode.The flow process of figure six is reboiler E111, the water cooler E116 of the overhead product of high pressure deethanizing column D102 through demethanizing tower D101.Enter knockout drum F102 behind the intermediate reboiler E112 of demethanizing tower and the water cooler E117, the section port gas of F102 outlet liquid through having reduced temperature and F102 after the vacuum flashing carries out entering demethanizing tower after the heat exchange in E121.Better choice is that the liquid that F102 exports is divided into two, and partially liq directly enters demethanizing tower, and the section port gas of partially liq and F102 carries out entering demethanizing tower again after the heat exchange.F102 exit gas without heat exchange goes to mix with the gas of process E121 heat exchange after the E105 heat exchange again, enters F103 after further cooling off and carry out gas-liquid separation in E106.The liquid of F103 outlet is carrying out removing demethanizing tower after the heat exchange in E122 with part F103 exit gas after the decompression.According to the requirement of A point control ethylene loss and reduction D101 rectifying section load, the temperature of decision D103 inlet gas.Further reduce the temperature of D103 inlet gas as need, D103 can be exported liquid and also in E122, go demethanizing tower after the heat exchange again as charging.Tower top temperature was lower than-101 ℃ when demethanizing tower was under low pressure operated, thereby can not directly remove to cool off overhead gas with the ethene refrigerant, but the heat exchange of overhead gas elder generation produces phlegma with ethene refrigerant and overhead gas cooling again.A phlegma part is as refluxing; A part is removed methane absorber D103.D103 outlet liquid returns D101 as first burst of charging.The still liquid of demethanizing tower removes ethylene rectification tower after acetylene hydrogenation reactor removes acetylene.
When charging has been forced into about 3MPa and has removed acetylene, when for example charging is from the cat head of the predepropanization tower of ethylene unit, then charging needn't be through overdraft (C101) and dehydration (V101), directly go to enter the high pressure deethanizing column after the E101 cooling, and the still liquid of demethanizing tower directly removes ethylene rectification tower without acetylene hydrogenation reactor.
The effect that the present invention reached:
One, improved the ethylene recovery rate
1. greatly reduced the ethylene loss that A is ordered
Because absorption tower D103 has a plurality of equilibrium separation levels and absorption agent, under the D103 Outlet Gas Temperature situation high than the F104 temperature out of flow process one or other similar flow processs, the ethylene loss that ethylene loss of the present invention is ordered than the F104 outlet A of an equilibrium separation level reduces more than 70%.
2. eliminated the ethylene loss that C is ordered in other flow processs
The liquid phase methane that flow process one or other similar flow processs are extracted out from the demethanizing tower uses as cryogen, vaporization back discharge system, wherein contained ethene complete loss.The liquid methane that the present invention extracts out uses as absorption agent, returns demethanizing tower behind the absorption tower, and therefore, no matter extract how many liquid phase methane and its ethylene concentration out, wherein contained ethene does not lose.
The ethylene loss of demethanizing therefore of the present invention system can reduce about 3/4ths.
Two, reduced the load of demethanizing tower
Because the temperature out of methane absorber D103 is than the F104 height of flow process one, methane content under feed composition and all identical condition of pressure in the exit gas is than flow process one height, therefore reduced the methane content that enters demethanizing tower, the ratio that reduces is relevant with the processing parameter on absorption tower, the inlet gas temperature that wherein most important parameter is D103, the ratio that the charging methane content reduces can reach 30～50%.When methane absorber is got along well previously described high pressure deethanizing column when combining, as hereinafter described flow process device three of embodiment and flow process four, reduce the methane content that advances demethanizing tower and can not reduce the load of demethanizing tower to significant degree, be the gas-liquid phase flow rate and the condenser of demethanizing tower, the load of reboiler can not reduce (table four column data sees below) by a relatively large margin, this is not cut because advance the c 3 fraction and/or part carbon two fractions of tower, it is basic identical that still liquid removes the requirement and the flow process 1 of the required tower still of light constituent rising quantity of steam, therefore the load of overhead condenser depends on the heat balance of full tower, even cat head methane load reduces, still can not reduce the thermal load of condenser.
The charging of demethanizing tower of the present invention has only methane and part carbon two fractions, and this moment, stripping section removed the required rising quantity of steam of light constituent methane in the still liquid than flow process one decline, for high pressure demethanizer, had descended about 40%; And for low pressure demethanizer, then descended about 60%, at this moment, no matter be high pressure demethanizing or low-pressure methane removing, full tower thermal equilibrium has not been the controlling factor of condenser duty, and cat head methane load reduces, and the required quantity of reflux of rectifying section just reduces, the gas-liquid phase load and the condenser duty that are rectifying section also reduce, and therefore the full tower load of demethanizing tower of the present invention descends.
Three, reduced energy expenditure
As mentioned before, flow process one and similar flow process thereof will reduce ethylene loss must increase energy expenditure, but not only the present invention when improving yield of ethene, do not increase energy consumption, reduced energy consumption on the contrary, reason is:
1. to reduce the measure of ethylene loss be to adopt absorption tower D103 in the present invention, remove to absorb ethene in the exit gas with methane, the liquid phase methane of extracting demethanizer column overhead out does not increase the load of demethanizing tower condenser as absorption agent, reason has 2 points: first, the liquid phase methane of extracting out returns demethanizing tower as first burst of charging, its effect is equivalent to reflux, and the then required quantity of reflux of absorption dose increase reduces.The secondth, D103 inlet gas temperature is than the F104 inlet gas temperature height of flow process one and similar flow process thereof, and the thermal load of E107 reduces, thereby reduces or cancelled the demand of E107 to liquid phase methane.Comprehensive these 2 reasons make the load of demethanizing tower condenser increase unlike flow process one and similar flow process thereof.
2. owing to preamble two, reduced the described reason of this joint of load of demethanizing tower, the condenser duty of demethanizing tower of the present invention is lower than other flow processs.
3. ethylene rectification tower of the present invention is formed charging inequality by two strands, and is energy-conservation than the ethylene rectification tower of sub-thread charging.The concentration of ethene is lower in the C-2-fraction, and energy-conservation effect is more remarkable, and the condenser duty of ethylene rectification tower can reduce by 4～10% after the sub-thread charging changes bifilar charging into.If increase the plate number again during enlarging, the superiority of bifilar charging is more remarkable.
Therefore the energy consumption of flow process two is lower than other flow processs.
Accompanying drawing and explanation thereof
Figure one is lighter hydrocarbons separation process one sketch
Figure two is technical process (flow process two) sketches that the present invention is applied to high pressure demethanizing system
Figure three is lighter hydrocarbons separation process three sketches
Figure four is lighter hydrocarbons separation process four sketches
Figure five is lighter hydrocarbons separation process five sketches
Figure six is technical process (flow process six) sketches that the present invention is applied to the low-pressure methane removing system
Device E115 D102 reflux cooler E116～E117 ethene refrigerant cooler E120 methane cooler E121～E122 heat exchanger F102～F105 knockout drum P101 domethanizing column reflux pump P102 D102 backflow booster R101 acetylene hydrogenation reactor V101 drier (R101A～R101B) press methane among the material code name explanation C3 carbon three above cut E4 ethene E6 ethane F charging H2 hydrogen LP low-pressure methane MP boils in device name explanation C101 feed compressor C102 methane compressor D101 domethanizing column D102 high pressure dethanizer D103 methane absorber D104 low pressure dethanizer D105 ethylene rectifying column E101 dryer feed cooler E102 drier aftercooler E103～E107 domethanizing column feed cooler E109 methane hydrogen cooler E110 domethanizing column condenser E111 domethanizing column reboiler E112 domethanizing column
Embodiment
Certain 450,000 tons of/year ethylene unit is intended enlarging to 600,000 tons/year, and is existing with identical design basis, with strict several different flow schemes, the relatively load of energy expenditure, ethylene loss and several main rectifying tower of these schemes of calculating of flowsheeting spare.The working pressure of demethanizing tower is 3.1MPA, and the load of the device that boils in boil in the demethanizing tower device and the ethylene column is to determine according to the integrated needs of heat.
Several different flow schemes are:
1. scheme one: the flow process one that preceding texts and pictures one are represented
2. scheme two: the flow process two that preceding texts and pictures two are represented, i.e. the present invention
3. scheme three: front-end deethanization and front-end hydrogenation flow process, see figure three
The difference of flow process three and flow process two is that mainly the D102 overhead product of flow process three has comprised whole carbon two fractions, and flow process two has only comprised about 50% carbon two fractions, and the difference that causes thus is:
(1). the D102 tower top temperature of flow process three is higher than flow process two, the cold that under the prerequisite that suitably improves the overhead product propylene content, can provide with junior propylene refrigerant, therefore saved the required ethene cold of D102 of flow process two, deliver to D102 after the propylene refrigerant cooling of phlegma through highest ranking with the D104 cat head, promptly the condenser of D104 provides the backflow of D102 and D104 simultaneously.
(2) overhead product of .D102 has removed C3 fraction basically, and therefore when it entered the cooling of demethanizing charging cooling system, the condensation of C-2-fraction all need provide cold by the ethene refrigerant.Because flow process three enters the C-2-fraction amount of demethanizing system and almost doubles than flow process two, double so required ethene cold is also approximate.Therefore, even the cold that D102 itself does not need ethene to provide, the power of ethylene refrigerator is still greater than flow process two.
(3). the load flow process two of demethanizing tower is less than flow process three.
(4). the ethylene column of flow process two is formed different chargings by two strands, and flow process three has only one charging, so the former ethylene column load is lower than the latter.
In addition, boil in the acetylene hydrogenation reactor position of flow process two and flow process three and the demethanizing tower hot integration mode of device is also different.
4. scheme four: see the flow process four that figure four is represented
Flow process four is that flow process one is made local modification, adopts the methane absorber D103 in the flow process two, and the exit gas that is about to E107 is delivered to D103, uses the liquid phase methane adsorption ethene wherein from the D301 return tank, and all the other are identical with flow process one.
5. scheme five: see the flow process five that figure five is represented
This flow process is the described double tower front-end deethanization of ZL92100471.0 flow process, only is to have increased high pressure front-end deethanization tower D102 with the difference of flow process one.
The calculation result of above-mentioned five kinds of schemes sees Table one to table four.
Can draw following results by table one to table four column data:
1. flow process two and flow process three all are double tower front-end deethanization flow process, but both compare, and three acc powers of flow process two hang down 3643kw, the condenser duty of propylene compressor, i.e. and the consumption of water coolant has reduced by 19.4 * 106KJ/h.
2. flow process four is compared with flow process one, and ethylene loss has reduced 1064T/Y, has reflected methane absorber improves the ethylene recovery rate in flow process four effect.
3. flow process five is compared with flow process one, and three acc powers have reduced 1611kw, has reflected that the cat head discharging only contains the energy-conserving action of high pressure front-end deethanization in flow process five of part C-2-fraction.
4. flow process two is compared with flow process one, not only three acc powers have reduced 2000kw, straight greater than flow process five with the difference of flow process one, ethylene loss has reduced 1106T/Y, greater than the difference of flow process four with flow process one, and the gas phase of demethanizing tower load descended 35～40%, and the liquid phase load has descended 33～55%.This advantage is the investment that has reduced deep cooling demethanizing system for new device, and for the enlarging of ethylene unit, is not only and has reduced investment, also has the retrofit work amount that reduces existing device, shortens remarkable advantages such as downtime period.
5. by the D104 still liquid data of table one as can be known, the propylene rate of recovery of flow process three is minimum, compares with other flow process, and propone output has reduced about 2000T/Y, this be because the tower top temperature of D102 higher due to.
When the working pressure of demethanizing tower during at 0.6～1.5MPa, the ethylene recovery rate between each flow process and the difference of energy consumption slightly change, but do not change relative size, and not changing flow process two is wherein to possess ethylene recovery rate height simultaneously, energy consumption is low, the conclusion of the flow process of three kinds of advantages of cryogenic system reduced investment.
Table one material is formed Kmol/h
| Charging | D104 still liquid | |||||
| Flow process one | Flow process two | Flow process three | Flow process four | Flow process five | ||
| H2 | ????1225.29 | |||||
| CO | ????16.24 | |||||
| CH4 | ????1982.43 | |||||
| C2H2 | ????34.04 | |||||
| C2H4 | ????2713.31 | |||||
| C2H6 | ????723.40 | ????0.22 | ??0.22 | ??0.55 | ????0.22 | ??0.22 |
| C3H4 | ????34.02 | ????34.02 | ??34.02 | ??34.00 | ????34.05 | ??34.02 |
| C3H6 | ????850.82 | ????846.24 | ??846.27 | ??840.24 | ????846.24 | ??846.27 |
| C3H8 | ????25.20 | ????25.18 | ??25.19 | ??25.10 | ????25.18 | ??25.19 |
| C4H6 | ????76.86 | ????76.86 | ??76.86 | ??76.86 | ????76.86 | ??76.86 |
| C4H8 | ????94.86 | ????94.86 | ??94.86 | ??94.86 | ????94.86 | ??94.86 |
| C4H10 | ????2.48 | ????2.48 | ??2.48 | ??2.48 | ????2.48 | ??2.48 |
| C5H12 | ????13.38 | ????13.38 | ??13.38 | ??13.38 | ????13.38 | ??13.38 |
| C6H6 | ????8.22 | ????8.22 | ??8.22 | ??8.22 | ????8.22 | ??8.22 |
| Add up to | ????7800.56 | ????1101.44 | ??1101.50 | ??1095.69 | ????1101.44 | ??1101.50 |
Table divinyl loss Kmol/h
| Flow process one | Flow process two | Flow process three | Flow process four | Flow process five | |
| The A point | ????3.5551 | ????0.5893 | ????0.2924 | ????0.4371 | ????3.2230 |
| The B point | ????1.3517 | ????0.6966 | ????1.2992 | ????1.1766 | ????1.0689 |
| The C point | ????1.4470 | ????0 | ????0 | ????0 | ????2.4369 |
| Total losses | ????6.3538 | ????1.2857 | ????1.5916 | ????1.6137 | ????6.7288 |
| The total losses difference | |||||
| Kmol/h | Benchmark | ????-5.0681 | ????-4.7622 | ????-4.740 | ????0.375 |
| T/Y | Benchmark | ????-1137 | ????-1069 | ????-1064 | ????84 |
Table three energy consumption relatively
| Flow process one | Flow process two | Flow process three | Flow process four | Flow process five | |
| Ethylene compressor KW/h | ??3923 | ??3959 | ??5278 | ??3910 | ????4076 |
| Propylene compressor KW/h | ??24325 | ??22485 | ??24751 | ??24330 | ????22416 |
| Low-pressure methane compressor KW/h | ??746 | ??550 | ??508 | ??499 | ????891 |
| Three machine total power KW/h | ??28994 | ??26994 | ??30757 | ??28739 | ????27383 |
| Three machine total power difference KW/h | Benchmark | ??-2000 | ??1543 | ??-255 | ????-1611 |
| Propylene condenser load 10 6KJ/h | ??231.26 | ??214.62 | ??234.01 | ??231.57 | ????214.52 |
| Load difference 10 6KJ/h | Benchmark | ??-16.64 | ??2.75 | ??-0.31 | ????-16.74 |
Table four rectifying tower duty ratio
| Flow process one | Flow process two | Flow process three | Flow process four | Flow process five | |
| Demethanizing tower D101 | |||||
| Cat head gas phase V 2?Kmol/h | ????3119 | ????2055 | ????3020 | ????3042 | ????2619 |
| Cat head liquid phase L 2?Kmol/h | ????1467 | ????1010 | ????1557 | ????1713 | ????1021 |
| Tower spirit phase V n-1?Kmol/h | ????2896 | ????1730 | ????3509 | ????2856 | ????1730 |
| Liquid phase L at the bottom of the tower n-1?Kmol/h | ????7394 | ????3223 | ????6995 | ????7357 | ????3221 |
| Condenser duty Q d10 6KJ/h | ????8.152 | ????6.218 | ????7.939 | ????8.613 | ????7.084 |
| Low pressure deethanizing column D104 | |||||
| Cat head gas phase V 2?Kmol/h | ????6195 | ????4133 | ????3474 | ????6196 | ????4133 |
| Cat head liquid phase L 2?Kmol/h | ????2379 | ????2166 | ????1941 | ????2739 | ????2167 |
| Tower spirit phase V n-1?Kmol/h | ????3276 | ????2398 | ????2772 | ????3272 | ????2398 |
| Liquid phase L at the bottom of the tower n-1?Kmol/h | ????4302 | ????3448 | ????3804 | ????4297 | ????3448 |
| Condenser duty Q d10 6KJ/h | ????20.290 | ????14.893 | ????29.832 | ????20.294 | ????14.893 |
| Ethylene column D105 | |||||
| Cat head gas phase V 2?Kmol/h | ????12817 | ????11979 | ????12344 | ????12822 | ????11957 |
| Cat head liquid phase L 2?Kmol/h | ????13054 | ????12196 | ????12544 | ????13058 | ????12176 |
| Tower spirit phase V n-1?Kmol/h | ????5551 | ????5684 | ????5505 | ????5554 | ????5672 |
| Liquid phase L at the bottom of the tower n-1?Kmol/h | ????6316 | ????6449 | ????6270 | ????6318 | ????6437 |
| Condenser duty Q d10 6KJ/h | ????119.31 | ????111.35 | ????116.33 | ????119.35 | ????111.17 |
Claims (1)
1. the light hydrocarbon separating method of one kind high ethylene recovery rate, energy-conservation and low demethanizing tower load, technical process is that lighter hydrocarbons are through overdraft, enter knockout drum after the cooling and carry out gas-liquid separation, liquid and gas enter the non-clear cutting that high pressure deethanizing column (D102) carries out carbon two fractions respectively.High pressure deethanizing column (D102) bottoms goes low pressure deethanizing column (D104) to carry out separating of carbon two and c 3 fraction, high pressure deethanizing column (D102) overhead product obtains through cooling and flash distillation step by step that liquid enters demethanizing tower (D101) as charging more than two strands.The gas of gas-liquid separator top exit is through further entering methane absorber (D103) after the cooling,, as absorption agent 99.5% ethene in the knockout drum exit gas absorbed with liquid phase methane.Methane absorber (D103) outlet at bottom liquid returns demethanizing tower as the top charging.Still liquid removes acetylene hydrogenation reactor and/or ethylene rectification tower (D105) and obtains ethene and ethane at the bottom of demethanizing tower (D101) tower.The feature of flow process is to have the cat head discharging only to contain the high pressure deethanizing column of part carbon two fractions in the upstream of demethanizing tower and is the absorption tower of absorption agent with methane.
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| DE3802553C2 (en) * | 1988-01-28 | 1996-06-20 | Linde Ag | Process for the separation of hydrocarbons |
| US4900347A (en) * | 1989-04-05 | 1990-02-13 | Mobil Corporation | Cryogenic separation of gaseous mixtures |
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