CN104098441A - Technology and device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gases and producing ethylene glycol through dimethyl oxalate hydrogenation - Google Patents
Technology and device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gases and producing ethylene glycol through dimethyl oxalate hydrogenation Download PDFInfo
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- CN104098441A CN104098441A CN201410353158.7A CN201410353158A CN104098441A CN 104098441 A CN104098441 A CN 104098441A CN 201410353158 A CN201410353158 A CN 201410353158A CN 104098441 A CN104098441 A CN 104098441A
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- LYCAIKOWRPUZTN-UHFFFAOYSA-N Ethylene glycol Chemical compound OCCO LYCAIKOWRPUZTN-UHFFFAOYSA-N 0.000 title claims abstract description 334
- 239000007789 gas Substances 0.000 title claims abstract description 277
- 238000005984 hydrogenation reaction Methods 0.000 title claims abstract description 250
- 238000005810 carbonylation reaction Methods 0.000 title claims abstract description 236
- 230000006315 carbonylation Effects 0.000 title claims abstract description 188
- LOMVENUNSWAXEN-UHFFFAOYSA-N Methyl oxalate Chemical compound COC(=O)C(=O)OC LOMVENUNSWAXEN-UHFFFAOYSA-N 0.000 title claims abstract description 93
- 230000015572 biosynthetic process Effects 0.000 title claims abstract description 40
- 238000003786 synthesis reaction Methods 0.000 title claims abstract description 40
- 238000005516 engineering process Methods 0.000 title abstract description 11
- OKKJLVBELUTLKV-UHFFFAOYSA-N Methanol Chemical compound OC OKKJLVBELUTLKV-UHFFFAOYSA-N 0.000 claims abstract description 900
- 238000011084 recovery Methods 0.000 claims abstract description 223
- 239000007788 liquid Substances 0.000 claims abstract description 212
- 238000005886 esterification reaction Methods 0.000 claims abstract description 122
- 238000000926 separation method Methods 0.000 claims abstract description 95
- BLLFVUPNHCTMSV-UHFFFAOYSA-N methyl nitrite Chemical compound CON=O BLLFVUPNHCTMSV-UHFFFAOYSA-N 0.000 claims abstract description 74
- 239000002994 raw material Substances 0.000 claims abstract description 64
- GRYLNZFGIOXLOG-UHFFFAOYSA-N Nitric acid Chemical compound O[N+]([O-])=O GRYLNZFGIOXLOG-UHFFFAOYSA-N 0.000 claims abstract description 59
- 229910017604 nitric acid Inorganic materials 0.000 claims abstract description 58
- 238000006243 chemical reaction Methods 0.000 claims abstract description 50
- 238000010926 purge Methods 0.000 claims abstract description 34
- 239000002699 waste material Substances 0.000 claims abstract description 33
- 238000004064 recycling Methods 0.000 claims abstract description 27
- 239000002253 acid Substances 0.000 claims abstract description 26
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 22
- 239000001257 hydrogen Substances 0.000 claims abstract description 20
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 20
- 230000008878 coupling Effects 0.000 claims abstract description 12
- 238000010168 coupling process Methods 0.000 claims abstract description 12
- 238000005859 coupling reaction Methods 0.000 claims abstract description 12
- 239000003507 refrigerant Substances 0.000 claims description 88
- 239000003054 catalyst Substances 0.000 claims description 86
- 239000012528 membrane Substances 0.000 claims description 73
- 239000000047 product Substances 0.000 claims description 72
- 238000000034 method Methods 0.000 claims description 69
- 238000005406 washing Methods 0.000 claims description 61
- 230000008569 process Effects 0.000 claims description 55
- 239000000463 material Substances 0.000 claims description 54
- 239000012071 phase Substances 0.000 claims description 50
- 239000000203 mixture Substances 0.000 claims description 44
- 239000007791 liquid phase Substances 0.000 claims description 40
- 238000012856 packing Methods 0.000 claims description 34
- 230000018044 dehydration Effects 0.000 claims description 33
- 238000006297 dehydration reaction Methods 0.000 claims description 33
- 238000001179 sorption measurement Methods 0.000 claims description 33
- 239000002808 molecular sieve Substances 0.000 claims description 30
- 238000010992 reflux Methods 0.000 claims description 30
- URGAHOPLAPQHLN-UHFFFAOYSA-N sodium aluminosilicate Chemical compound [Na+].[Al+3].[O-][Si]([O-])=O.[O-][Si]([O-])=O URGAHOPLAPQHLN-UHFFFAOYSA-N 0.000 claims description 30
- LFQSCWFLJHTTHZ-UHFFFAOYSA-N Ethanol Chemical compound CCO LFQSCWFLJHTTHZ-UHFFFAOYSA-N 0.000 claims description 26
- 238000010521 absorption reaction Methods 0.000 claims description 18
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 claims description 18
- 230000032050 esterification Effects 0.000 claims description 16
- 239000007809 chemical reaction catalyst Substances 0.000 claims description 14
- 238000012545 processing Methods 0.000 claims description 14
- 239000012510 hollow fiber Substances 0.000 claims description 13
- 238000007599 discharging Methods 0.000 claims description 10
- 239000003463 adsorbent Substances 0.000 claims description 9
- 239000002131 composite material Substances 0.000 claims description 9
- 229920006395 saturated elastomer Polymers 0.000 claims description 8
- 239000007795 chemical reaction product Substances 0.000 claims description 7
- 238000001035 drying Methods 0.000 claims description 6
- 230000008929 regeneration Effects 0.000 claims description 6
- 238000011069 regeneration method Methods 0.000 claims description 6
- 238000000746 purification Methods 0.000 claims description 5
- 230000002378 acidificating effect Effects 0.000 claims description 4
- BRPQOXSCLDDYGP-UHFFFAOYSA-N calcium oxide Chemical compound [O-2].[Ca+2] BRPQOXSCLDDYGP-UHFFFAOYSA-N 0.000 claims description 3
- 239000000292 calcium oxide Substances 0.000 claims description 3
- ODINCKMPIJJUCX-UHFFFAOYSA-N calcium oxide Inorganic materials [Ca]=O ODINCKMPIJJUCX-UHFFFAOYSA-N 0.000 claims description 3
- 230000001476 alcoholic effect Effects 0.000 claims description 2
- 125000004122 cyclic group Chemical group 0.000 claims description 2
- 230000001172 regenerating effect Effects 0.000 claims 1
- IEJIGPNLZYLLBP-UHFFFAOYSA-N dimethyl carbonate Chemical compound COC(=O)OC IEJIGPNLZYLLBP-UHFFFAOYSA-N 0.000 abstract description 23
- 239000002912 waste gas Substances 0.000 abstract description 6
- 230000000694 effects Effects 0.000 abstract description 4
- 150000007513 acids Chemical class 0.000 abstract 2
- 238000004519 manufacturing process Methods 0.000 description 12
- 238000005265 energy consumption Methods 0.000 description 8
- 229910002091 carbon monoxide Inorganic materials 0.000 description 7
- VNWKTOKETHGBQD-UHFFFAOYSA-N methane Chemical group C VNWKTOKETHGBQD-UHFFFAOYSA-N 0.000 description 7
- IJGRMHOSHXDMSA-UHFFFAOYSA-N Atomic nitrogen Chemical compound N#N IJGRMHOSHXDMSA-UHFFFAOYSA-N 0.000 description 6
- 239000000945 filler Substances 0.000 description 6
- 239000000243 solution Substances 0.000 description 6
- MWUXSHHQAYIFBG-UHFFFAOYSA-N Nitric oxide Chemical compound O=[N] MWUXSHHQAYIFBG-UHFFFAOYSA-N 0.000 description 5
- 239000000835 fiber Substances 0.000 description 5
- 229910052757 nitrogen Inorganic materials 0.000 description 5
- JVTAAEKCZFNVCJ-UHFFFAOYSA-M Lactate Chemical compound CC(O)C([O-])=O JVTAAEKCZFNVCJ-UHFFFAOYSA-M 0.000 description 4
- 239000003245 coal Substances 0.000 description 4
- 238000011049 filling Methods 0.000 description 4
- 238000002156 mixing Methods 0.000 description 4
- 239000002351 wastewater Substances 0.000 description 4
- 230000007613 environmental effect Effects 0.000 description 3
- 239000000126 substance Substances 0.000 description 3
- HZAXFHJVJLSVMW-UHFFFAOYSA-N 2-Aminoethan-1-ol Chemical compound NCCO HZAXFHJVJLSVMW-UHFFFAOYSA-N 0.000 description 2
- UGFAIRIUMAVXCW-UHFFFAOYSA-N Carbon monoxide Chemical compound [O+]#[C-] UGFAIRIUMAVXCW-UHFFFAOYSA-N 0.000 description 2
- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 description 2
- 239000005977 Ethylene Substances 0.000 description 2
- 230000008901 benefit Effects 0.000 description 2
- 239000006227 byproduct Substances 0.000 description 2
- 238000009903 catalytic hydrogenation reaction Methods 0.000 description 2
- 239000012043 crude product Substances 0.000 description 2
- 238000004134 energy conservation Methods 0.000 description 2
- 238000002474 experimental method Methods 0.000 description 2
- 238000000605 extraction Methods 0.000 description 2
- XLYOFNOQVPJJNP-ZSJDYOACSA-N heavy water Substances [2H]O[2H] XLYOFNOQVPJJNP-ZSJDYOACSA-N 0.000 description 2
- CBOIHMRHGLHBPB-UHFFFAOYSA-N hydroxymethyl Chemical compound O[CH2] CBOIHMRHGLHBPB-UHFFFAOYSA-N 0.000 description 2
- 230000004048 modification Effects 0.000 description 2
- 238000012986 modification Methods 0.000 description 2
- 229920000728 polyester Polymers 0.000 description 2
- 230000009467 reduction Effects 0.000 description 2
- 239000013589 supplement Substances 0.000 description 2
- 230000002194 synthesizing effect Effects 0.000 description 2
- 238000012546 transfer Methods 0.000 description 2
- 241000669618 Nothes Species 0.000 description 1
- 238000004458 analytical method Methods 0.000 description 1
- 230000002528 anti-freeze Effects 0.000 description 1
- 239000012736 aqueous medium Substances 0.000 description 1
- 230000009286 beneficial effect Effects 0.000 description 1
- 238000009835 boiling Methods 0.000 description 1
- 230000003197 catalytic effect Effects 0.000 description 1
- 230000008859 change Effects 0.000 description 1
- 230000006835 compression Effects 0.000 description 1
- 238000007906 compression Methods 0.000 description 1
- 238000009833 condensation Methods 0.000 description 1
- 230000005494 condensation Effects 0.000 description 1
- 239000002826 coolant Substances 0.000 description 1
- 239000000498 cooling water Substances 0.000 description 1
- 230000007547 defect Effects 0.000 description 1
- 238000006356 dehydrogenation reaction Methods 0.000 description 1
- 238000013461 design Methods 0.000 description 1
- 238000011161 development Methods 0.000 description 1
- 230000018109 developmental process Effects 0.000 description 1
- 238000009792 diffusion process Methods 0.000 description 1
- 238000004090 dissolution Methods 0.000 description 1
- 238000004821 distillation Methods 0.000 description 1
- 238000009826 distribution Methods 0.000 description 1
- 230000009977 dual effect Effects 0.000 description 1
- 239000002360 explosive Substances 0.000 description 1
- 230000002349 favourable effect Effects 0.000 description 1
- 239000012467 final product Substances 0.000 description 1
- 238000010438 heat treatment Methods 0.000 description 1
- 230000036571 hydration Effects 0.000 description 1
- 238000006703 hydration reaction Methods 0.000 description 1
- 150000002431 hydrogen Chemical class 0.000 description 1
- 238000009776 industrial production Methods 0.000 description 1
- 239000011261 inert gas Substances 0.000 description 1
- 238000011031 large-scale manufacturing process Methods 0.000 description 1
- 239000000314 lubricant Substances 0.000 description 1
- 239000002609 medium Substances 0.000 description 1
- 125000002496 methyl group Chemical group [H]C([H])([H])* 0.000 description 1
- 239000011259 mixed solution Substances 0.000 description 1
- 238000005457 optimization Methods 0.000 description 1
- 230000003647 oxidation Effects 0.000 description 1
- 238000007254 oxidation reaction Methods 0.000 description 1
- 230000035699 permeability Effects 0.000 description 1
- 239000003208 petroleum Substances 0.000 description 1
- 239000004014 plasticizer Substances 0.000 description 1
- 231100000614 poison Toxicity 0.000 description 1
- 229920000642 polymer Polymers 0.000 description 1
- 239000002861 polymer material Substances 0.000 description 1
- 229920005903 polyol mixture Polymers 0.000 description 1
- 238000002360 preparation method Methods 0.000 description 1
- 150000003254 radicals Chemical class 0.000 description 1
- 239000002904 solvent Substances 0.000 description 1
- 239000003440 toxic substance Substances 0.000 description 1
- 239000002918 waste heat Substances 0.000 description 1
Classifications
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
- C07C29/132—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group
- C07C29/136—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH
- C07C29/147—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof
- C07C29/149—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof with hydrogen or hydrogen-containing gases
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C201/00—Preparation of esters of nitric or nitrous acid or of compounds containing nitro or nitroso groups bound to a carbon skeleton
- C07C201/04—Preparation of esters of nitrous acid
-
- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C67/00—Preparation of carboxylic acid esters
- C07C67/36—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/10—Process efficiency
- Y02P20/129—Energy recovery, e.g. by cogeneration, H2recovery or pressure recovery turbines
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- Chemical & Material Sciences (AREA)
- Organic Chemistry (AREA)
- Chemical Kinetics & Catalysis (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The invention relates to a technology and a device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gases and producing ethylene glycol through dimethyl oxalate hydrogenation. The technology comprises the following steps: adopting industrial NO, O2 and methanol as raw materials for an esterification reaction to produce methyl nitrite; adopting industrial CO and methyl nitrite for a carbonylation reaction in a plate reactor to produce carbonylation products, which mainly include dimethyl oxalate and dimethyl carbonate; separating the carbonylation products to obtain dimethyl carbonate products; subsequently adding hydrogen into dimethyl oxalate in the plate reactor to produce ethylene glycol products; conducting the coupling recovery treatment on waste acids in the esterification reaction and purge gases in the carbonylation reaction for recycling. The device system comprises an esterification reaction system, a carbonylation reaction system, a coupling recovery system for purge gases and waste acids and a hydrogenation reaction system. The technology has the characteristic that device consumption is remarkably reduced, and particularly the nitric acid waste liquid recycling and purge gas recycling technologies as well as the separation technologies thereof are highly coupled; recycling of the raw materials in reaction waste gases is realized, and the effect is remarkable.
Description
Technical Field
The invention relates to a process and a device system for preparing ethylene glycol from industrial synthesis gas, in particular to a process and a device system for preparing ethylene glycol from dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and hydrogenation.
Background
Ethylene glycol is a chemical with wide application, is mainly applied to various production fields such as polyester fibers (PET), antifreeze, ethanolamine, explosives and the like, is used as a solvent, a lubricant and a plasticizer in a large amount, and is approximately 95% applied to the PET polyester industry. Currently, ethylene glycol production in industry mainly adopts a route of preparing ethylene glycol by gas phase oxidation of petroleum ethylene and then liquid phase catalytic hydration. However, with the international oil price being high for a long time in recent years, the industrial chain for preparing ethylene glycol by using ethylene as a raw material faces huge pressure in the world at present. Therefore, the technical route for preparing ethylene glycol by using synthesis gas has attracted more and more attention due to the low production cost.
At present, a shell and tube reactor is mainly adopted in the technical process of preparing ethylene glycol from coal, and the common problems of low reaction heat transfer efficiency and low utilization coefficient and filling coefficient of a catalyst influence the production capacity of the reactor.
The patent (publication No. CN101462961) provides a process for producing ethylene glycol and CO-producing dimethyl carbonate, the process comprises a process of synthesizing dimethyl oxalate and dimethyl carbonate from CO and methyl nitrite, a process of obtaining dimethyl carbonate products through distillation and separation, a process of synthesizing ethylene glycol through catalytic hydrogenation of heavy component dimethyl oxalate, and a regeneration reaction process of methyl nitrite in a system. However, the reactor adopts a tubular reactor, waste gas and waste liquid generated in the reaction process are not recycled, the energy consumption of the device is high, and the increasing environmental protection requirements of the state cannot be met.
The patent (publication No. CN101830806) discloses a method and a device for coproducing dimethyl carbonate and dimethyl oxalate, wherein the patent adopts two carbonylation reactors, the first is a dimethyl carbonate synthesis reactor, the second is a dimethyl oxalate reactor, methyl nitrite is generated after reaction and then respectively enters the two reactors to respectively generate dimethyl carbonate and dimethyl oxalate, and then products are separated and purified. The patent does not optimize the energy of the whole process flow, and does not disclose environmental protection measures necessary in the reaction process. The experimental procedure has not been an industrial process.
In addition, the loss of NO during the purge and the disposal of nitric acid by-products generated during the reaction are problematic. Patent CN201210531022.1 discloses a process in which the nitric acid produced is concentrated and then reacted with a portion of the NO-containing recycle gas to produce NO2And replenishing to return to a methyl nitrite regeneration reactor. However, the circulating gas containing NO also contains a large amount of gases such as methyl nitrite and methanol, which can also react with concentrated nitric acid, and the product is complex, thereby affecting the efficiency of the device.
In summary, the prior art for preparing ethylene glycol from coal mainly has the defects of low utilization rate of catalyst, low filling coefficient of catalyst, incapability of fully utilizing valuable gas in a device, environment pollution and incapability of fully utilizing system heat of the device, thereby causing unsatisfactory social and economic benefits.
Disclosure of Invention
The invention aims to solve the problems that the prior ethylene glycol production technology has low raw material utilization rate, high production cost, low catalyst utilization rate, low filling coefficient, overlarge equipment investment, single-series equipment cannot adapt to the upsizing of a device, the system consumption is high, the use of the device cannot meet the increasing requirements of China on the industrial environment and the like, and provides a process for improving the production capacity of the single-series device, treating tail gas, recycling and comprehensively utilizing raw materials and a device system thereof.
The invention is realized by the following technical scheme:
a device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation comprises a carbonylation reaction system, an esterification reaction system, a purge gas and waste acid coupling recovery system and a hydrogenation reaction system;
the carbonylation reaction system comprises a carbonylation reactor, a first gas-liquid separator, a methanol washing tower, a methanol rectifying tower and a DMO rectifying tower; the carbonylation reactor is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the first gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the methanol washing tower is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the methanol rectifying tower is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the DMO rectifying tower is provided with a lower feed inlet, a top outlet and a bottom outlet;
the esterification reaction system comprises an esterification reaction tower and a methanol recovery tower; the esterification reaction tower is provided with a top feed inlet, an upper feed inlet, a plurality of lower feed inlets, a middle reflux inlet, a top outlet and a bottom outlet; the methanol recovery tower is provided with a middle-lower part feed inlet, a top outlet and a bottom outlet;
the purge gas and waste acid coupling recovery system comprises a nitric acid concentration tower, an NO recovery tower, an MN recovery tower and a pressure swing adsorption tank; the nitric acid concentration tower is provided with a middle feeding hole, a top outlet and a bottom outlet; the NO recovery tower is provided with a top feed inlet, a middle feed inlet, a bottom feed inlet, a top outlet and a bottom outlet; the MN recovery tower is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the pressure swing adsorption tank is provided with a feed inlet, a recovered gas outlet and an exhaust gas outlet;
the hydrogenation reaction system comprises a hydrogenation circulating compressor, a hydrogenation reactor, a second gas-liquid separator, a membrane separator, a methanol separation tower, a light component rectifying tower and an ethylene glycol product tower; the hydrocycle compressor comprises an inlet and an outlet; the hydrogenation reactor is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the second gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the membrane separator is provided with a feed inlet, a recovered gas outlet and a discharged gas outlet; the methanol separation tower is provided with a middle feed inlet, a top noncondensable gas outlet, a top liquid phase light component outlet and a bottom liquid phase heavy component outlet; the light component rectifying tower is provided with a lower feed inlet, a top outlet and a bottom outlet; the ethylene glycol product tower is provided with a lower feed inlet, a top outlet, an upper outlet and a bottom outlet;
the top feed inlet of the carbonylation reactor is connected with a CO raw material pipeline and N2The raw material pipelines are connected through pipelines; a discharge hole at the bottom of the carbonylation reactor is connected with a feed inlet of the first gas-liquid separator through a pipeline; the gas outlet of the first gas-liquid separator is connected with the lower feed inlet of the methanol washing tower through a pipeline; a liquid outlet of the first gas-liquid separator is connected with an upper feed inlet of the methanol rectifying tower through a pipeline; the outlet at the top of the methanol washing tower is provided with a branch outlet A and a branch outlet B, the branch outlet A is connected with a lower feed inlet of the esterification reaction tower through a pipeline, and the branch outlet B is recovered with NOThe feed inlets at the bottom of the tower are connected by pipelines; the bottom outlet of the methanol washing tower is connected with the lower feed inlet of the methanol rectifying tower through a pipeline; the top outlet of the methanol rectifying tower is connected with the upper feed inlet of the esterification reaction tower through a pipeline; the bottom outlet of the methanol rectifying tower is connected with the lower feed inlet of the DMO rectifying tower through a pipeline; the bottom outlet of the DMO rectifying tower is connected with the top feed inlet of the hydrogenation reactor through a pipeline, and the top outlet of the DMO rectifying tower is a DMC product outlet;
other lower part feed inlets of the esterification reaction tower, an NO raw material pipeline and a multi-path O2The raw material pipelines are respectively connected through pipelines; a feed inlet at the top of the esterification reaction tower is connected with a methanol raw material pipeline through a pipeline; a branch outlet C and a branch outlet D are arranged at an outlet at the bottom of the esterification tower, the branch outlet C is connected with a middle reflux inlet of the esterification tower through a pipeline, and the branch outlet D is connected with a lower feeding port of the methanol recovery tower through a pipeline; the top outlet of the esterification reaction tower is connected with the top feed inlet of the carbonylation reactor through a pipeline; a branch outlet E and a branch outlet F are arranged at an outlet at the top of the methanol recovery tower, the branch outlet E is connected with an upper feed inlet of the esterification reaction tower through a pipeline, and the branch outlet F is connected with an upper feed inlet of the MN recovery tower through a pipeline; the bottom outlet of the methanol recovery tower is connected with the middle feed inlet of the nitric acid concentration tower through a pipeline;
an outlet at the top of the nitric acid concentration tower is a waste liquid outlet; the outlet at the bottom of the nitric acid concentration tower is connected with the middle feed inlet of the NO recovery tower through a pipeline; the top outlet of the NO recovery tower is connected with the lower feed inlet of the MN recovery tower through a pipeline; the bottom outlet of the NO recovery tower is connected with the feeding port at the middle lower part of the methanol recovery tower through a pipeline; the top outlet of the MN recovery tower is connected with the feed inlet of the pressure swing adsorption tank through a pipeline; the bottom outlet of the MN recovery tower is connected with the upper feed inlet of the esterification reaction tower through a pipeline; the recycled gas outlet of the pressure swing adsorption tank is connected with the top feed inlet of the carbonylation reactor through a pipeline; the exhaust gas outlet of the pressure swing adsorption tank is connected with an out-of-range recovery device through a pipeline;
the inlet of the hydrogenation circulating compressor is connected with an industrial hydrogen raw material pipeline through a pipeline, and the outlet of the hydrogenation circulating compressor is connected with the top feed inlet of the hydrogenation reactor through a pipeline; a discharge hole at the bottom of the hydrogenation reactor is connected with a feed inlet of the second gas-liquid separator through a pipeline; a gas outlet of the second gas-liquid separator is provided with a branch outlet G and a branch outlet H, the branch outlet G is connected with an inlet of the hydrogenation circulating compressor through a pipeline, and the branch outlet H is connected with a feed inlet of the membrane separator through a pipeline; the liquid outlet of the second gas-liquid separator is connected with the lower feed inlet of the methanol separation tower through a pipeline; the top noncondensable gas outlet of the methanol separation tower is connected with the feed inlet of the membrane separator through a pipeline; a top liquid phase light component outlet of the methanol separation tower is provided with a branch outlet I and a branch outlet J, the branch outlet I is connected with an upper feed inlet of the methanol washing tower through a pipeline, and the branch outlet J is connected with a top feed inlet of the NO recovery tower through a pipeline; a bottom liquid phase heavy component outlet of the methanol separation tower is connected with a lower feed inlet of the light component rectifying tower through a pipeline; a light component outlet at the top of the light component rectifying tower is connected with an out-of-range alcohol recovery device through a pipeline; a heavy component outlet at the bottom of the light component rectifying tower is connected with a feeding port at the lower part of the ethylene glycol product tower through a pipeline; the top outlet of the ethylene glycol product tower is connected with an out-of-range 1, 2-BDO recovery processing device through a pipeline; the bottom outlet of the ethylene glycol product tower is connected with an outside recovery processing device through a pipeline; an outlet at the upper part of the ethylene glycol product tower is an ethylene glycol product outlet; and the exhaust gas outlet of the membrane separator is connected with an out-of-range recovery device through a pipeline, and the recovered gas outlet of the membrane separator is connected with the top feed inlet of the hydrogenation reactor through a pipeline.
A dehydration tower is connected outside the carbonylation reactor; the dehydration tower is provided with a feed inlet and a dry gas outlet; the top outlet of the esterification reaction tower and the recovered gas outlet of the pressure swing adsorption tank are connected with the feed inlet of the dehydration tower through pipelines; and a dry gas outlet of the dehydration tower is connected with a feed inlet at the top of the carbonylation reactor through a pipeline.
The dehydration tower consists of two molecular sieve dryers A and B which alternately run and regenerate; the molecular sieve dryer A and the molecular sieve dryer B are filled with adsorbents; the adsorbent is selected from a 3A molecular sieve, a 4A molecular sieve, a 5A molecular sieve, a 9A molecular sieve and calcium oxide.
A discharge hole at the bottom of the carbonylation reactor is connected with an outlet heat exchanger I; the outlet heat exchanger I is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the CO raw material pipeline and N2A raw material pipeline and a dry gas outlet of the dehydration tower are connected with a cold material flow inlet of the outlet heat exchanger I through pipelines; a cold material flow outlet of the outlet heat exchanger I is connected with a top feed inlet of the carbonylation reactor through a pipeline; a discharge port at the bottom of the carbonylation reactor is connected with a hot material flow inlet of the outlet heat exchanger I through a pipeline; and the hot material outlet of the outlet heat exchanger I is connected with the feed inlet of the first gas-liquid separator through a pipeline.
A steam drum I is connected outside the carbonylation reactor; the steam pocket I is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a vapor outlet; a refrigerant inlet of the steam pocket I is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum I is connected with a refrigerant inlet at the bottom of the carbonylation reactor through a pipeline; a top refrigerant outlet of the carbonylation reactor is connected with a vapor-liquid mixture inlet of the steam drum I through a pipeline; and a steam outlet of the steam drum I is connected with an out-of-range steam recovery system through a pipeline.
A carbonylation circulating compressor is connected between the branch outlet A of the methanol washing tower and the lower feed inlet of the esterification reaction tower; the carbonylation circulating compressor is provided with an inlet and an outlet; the branch outlet A is connected with the inlet of the carbonylation circulating compressor through a pipeline; the outlet of the carbonylation circulating compressor is connected with the lower feed inlet of the esterification reaction tower through a pipeline.
A compressor is connected between the top outlet of the NO recovery tower and the bottom feed inlet of the MN recovery tower; the compressor is provided with an inlet and an outlet; the top outlet of the NO recovery tower is connected with the inlet of the compressor through a pipeline; the outlet of the compressor is connected with the bottom feed inlet of the MN recovery tower through a pipeline.
A discharge hole at the bottom of the hydrogenation reactor is connected with an outlet heat exchanger II; the outlet heat exchanger II is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the bottom outlet of the DMO rectifying tower, the recovered gas outlet of the membrane separator and the outlet of the hydrogenation circulating compressor are connected with the cold stream inlet of the outlet heat exchanger II through pipelines; a cold material flow outlet of the outlet heat exchanger II is connected with a top feeding hole of the hydrogenation reactor through a pipeline; a discharge hole at the bottom of the hydrogenation reactor is connected with a hot material flow inlet of the outlet heat exchanger II through a pipeline; and the hot material outlet of the outlet heat exchanger II is connected with the feed inlet of the second gas-liquid separator through a pipeline.
A top feeding hole of the hydrogenation reactor is connected with a start-up heater; the start-up heater is provided with a feed inlet and a discharge outlet; a cold material flow outlet of the outlet heat exchanger II is connected with a feed inlet of the start-up heater through a pipeline; and the discharge hole of the start-up heater is connected with the top feed inlet of the hydrogenation reactor through a pipeline.
A steam drum II is connected outside the hydrogenation reactor; the steam pocket II is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a vapor outlet; a refrigerant inlet of the steam pocket II is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum II is connected with a refrigerant inlet at the bottom of the hydrogenation reactor through a pipeline; a top refrigerant outlet of the hydrogenation reactor is connected with a vapor-liquid mixture inlet of the steam drum II through a pipeline; and a steam outlet of the steam drum II is connected with an out-of-range steam recovery system through a pipeline.
The second gas-liquid separator comprises a high-pressure gas-liquid separator and a low-pressure gas-liquid separator; the high-pressure gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the low-pressure gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the hot material outlet of the outlet heat exchanger II is connected with the feed inlet of the high-pressure gas-liquid separator through a pipeline; the gas outlet of the high-pressure gas-liquid separator is provided with a branch outlet K and a branch outlet L, the branch outlet K is connected with the inlet of the hydrogenation circulating compressor through a pipeline, and the branch outlet L is connected with the feed inlet of the low-pressure gas-liquid separator through a pipeline; a liquid outlet of the high-pressure gas-liquid separator is connected with a middle feed inlet of the methanol separation tower through a pipeline; the gas outlet of the low-pressure gas-liquid separator is connected with the feed inlet of the membrane separator through a pipeline; and a liquid outlet of the low-pressure gas-liquid separator is connected with a middle feed inlet of the methanol separation tower through a pipeline.
A methanol absorption tank is arranged in front of a feed inlet of the membrane separator; the methanol absorption tank is provided with a feed inlet and a purified gas outlet; a gas outlet of the low-pressure gas-liquid separator and a top noncondensable gas outlet of the methanol separation tower are connected with a feed inlet of the methanol absorption tank through pipelines; and a purified gas outlet of the methanol absorption tank is connected with a feed inlet of the membrane separator through a pipeline.
Preferably, the carbonylation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor.
Preferably, the carbonylation reactor is a plate type fixed bed carbonylation reactor.
Preferably, a plate group fixing cavity is arranged in the center of the plate type fixed bed carbonylation reactor, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed carbonylation reactor; the catalyst bed layer is filled with a carbonylation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate type fixed bed carbonylation reactor, a bottom refrigerant inlet of the plate type fixed bed carbonylation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate type fixed bed carbonylation reactor through a pipeline; and the top inlet of the plate fixed bed carbonylation reactor is connected with the top inlet of the catalyst bed layer through a pipeline, and the top outlet of the plate group fixing cavity is connected with the top refrigerant outlet of the plate fixed bed carbonylation reactor through a pipeline.
Preferably, the esterification reaction tower is a packed tower.
Preferably, the esterification reaction column is a tray-filler mixing column having both a tray portion and a filler-packed portion.
Preferably, the methanol washing tower, the methanol rectifying tower, the methanol recovery tower, the NO recovery tower, the MN recovery tower, the DMO rectifying tower and the nitric acid concentration tower are packed towers, plate towers or bubble cap towers.
Preferably, the filler filled in the packed tower is random packing or efficient regular packing; the random packing is in a saddle shape, a Raschig ring, a pall ring, a wheel shape, a rectangular saddle ring, a spherical shape or a columnar shape; the efficient structured packing is corrugated packing, grid packing or pulse packing.
Preferably, the hydrogenation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor.
More preferably, the hydrogenation reactor is a plate-type fixed bed hydrogenation reactor.
Preferably, a plate group fixing cavity is arranged in the center of the plate type fixed bed hydrogenation reactor, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed hydrogenation reactor; the catalyst bed layer is filled with a hydrogenation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate-type fixed bed hydrogenation reactor, a bottom refrigerant inlet of the plate-type fixed bed hydrogenation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate-type fixed bed hydrogenation reactor through a pipeline; the top of plate fixed bed hydrogenation ware, plate fixed bed hydrogenation ware's top feed inlet with the top entry pipe line connection of catalyst bed layer, the top export in the fixed chamber of plate group with plate fixed bed hydrogenation ware's top refrigerant export pipe line connection.
Preferably, the membrane separator is formed by connecting 1-100 hollow fiber membrane modules in parallel or in series.
A process for preparing ethanediol from dimethyl oxalate by high-pressure oxonation of industrial synthetic gas and hydrogenating it features use of industrial NO and O2And methanol are taken as raw materials to generate esterification reaction to generate methyl nitrite, then industrial grade CO and methyl nitrite are used for generating carbonylation products which mainly comprise dimethyl oxalate and dimethyl carbonate, the carbonylation products are separated to obtain dimethyl carbonate products, and the dimethyl oxalate is hydrogenated subsequently to generate ethylene glycol products; and the waste acid of the esterification reaction and the purge gas of the carbonylation reaction are recycled through coupling recovery treatment.
The reaction equation is as follows:
esterification reaction: 4NO + O2+4CH3OH→4CH3ONO+2H2O；
And (3) carbonylation reaction: 2CO +2CH3ONO→(COOCH3)2+2NO；
Hydrogenation reaction: (COOCH)3)2+4H2→(CH2OH)2+2CH3OH；
And (3) total reaction: 4CO + O2+8H2→2(CH2OH)2+2H2O；
The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation specifically comprises the following steps:
(1) introducing industrial grade NO and O into an esterification reaction tower2Carrying out esterification reaction with methanol; introducing the methyl nitrite mixed gas at the top of the esterification reaction tower into a carbonylation reactor for carbonylation reaction; refluxing part of the acidic alcohol solution in the tower kettle of the esterification reaction tower to the esterification reaction tower, and introducing part of the acidic alcohol solution into a methanol recovery tower; part of the methanol recovered from the top of the methanol recovery tower is recycled to the esterification reaction tower for recycling, and the rest of the methanol enters the MN recovery tower to be used as washing liquid; the waste acid at the tower kettle of the methanol recovery tower enters a nitric acid concentration tower for concentration treatment;
(2) methyl nitrite from esterification reaction tower and industrial grade CO and N2Feeding the feed into a carbonylation reactor, and carrying out carbonylation reaction in the presence of a carbonylation reaction catalyst; the temperature of the carbonylation reaction is 30-200 ℃, the reaction pressure is 1-10 MPa, and the gas hourly space velocity is 3000-30000 h-1；
(3) The carbonylation product enters a first gas-liquid separator to carry out gas-liquid separation, the gas phase enters a methanol washing tower, and the liquid phase enters a methanol rectifying tower; the gas phase component at the top of the methanol washing tower is partially circulated to the esterification reaction tower, and part of the gas phase component is used as purge gas to enter an NO recovery tower for recovery treatment; the liquid phase component of the tower bottom of the methanol washing tower enters a methanol rectifying tower for rectification and separation; the mixture of methanol and methyl nitrite recovered from the top of the methanol rectifying tower is recycled to the esterification reaction tower for reuse, and heavy components in the tower kettle enter the DMO rectifying tower; obtaining a DMC product at the top of the DMO rectifying tower, and enabling dimethyl oxalate components at the bottom of the rectifying tower to enter a hydrogenation reactor for hydrogenation reaction;
(4) concentrating waste acid from a methanol recovery tower by a nitric acid concentration tower until the concentration of nitric acid is 10-68 wt%, and circulating the waste acid to an NO recovery tower; in the NO recovery tower, concentrated nitric acid, methanol and purge gas from a methanol washing tower are subjected to esterification regeneration reaction; gas-phase light components at the top of the NO recovery tower enter an MN recovery tower, and the nitric acid waste liquid containing methanol generated at the bottom of the tower is circulated to the methanol recovery tower for further recovery treatment; in the MN recovery tower, gas phase feed is washed by recovered methanol and then enters a pressure swing adsorption tank, and the bottom of the MN recovery tower contains methyl nitriteThe alcohol solution circularly enters an esterification reaction tower; CO separated from pressure swing adsorption tank2Discharging to outside for treatment, recovering N2And the purified CO gas enters a carbonylation reactor for cyclic utilization;
(5) mixing dimethyl oxalate components from the tower bottom of the DMO rectifying tower with industrial hydrogen pressurized by a hydrogenation circulating compressor, then feeding the mixture into a hydrogenation reactor, and carrying out hydrogenation reaction in the presence of a hydrogenation catalyst to generate methanol, ethylene glycol and the like; the temperature of the hydrogenation reaction is 160-320 ℃, the reaction pressure is 1-10 MP, and the liquid hourly space velocity is 1-3 Kg/Kg.h;
(6) and the hydrogenated product enters a second gas-liquid separator for gas-liquid separation, a gas phase part is pressurized by the hydrogenation circulating compressor and then circulates to the hydrogenation reactor, a part of the gas phase enters a membrane separator for recycling and then returns to the hydrogenation reactor, and a liquid phase enters an ethylene glycol product tower for separation to obtain an ethylene glycol product.
Wherein,
preferably, the carbonylation reactor is externally connected with a dehydration tower; and the gas phase recovered by the pressure swing adsorption tank and the methyl nitrite mixed gas from the top of the esterification reaction tower enter a carbonylation reactor for carbonylation after moisture is removed by the dehydration tower.
Preferably, the dehydration tower consists of two molecular sieve dryers A and B which alternately run and regenerate; the molecular sieve dryer A and the molecular sieve dryer B are filled with adsorbents; the adsorbent is selected from a 3A molecular sieve, a 4A molecular sieve, a 5A molecular sieve, a 9A molecular sieve and calcium oxide. The operation temperature of the molecular sieve dryer A and the molecular sieve dryer B is 40-260 ℃, and the pressure is 1-10 MPa. All pressures in the present invention are indicated as gauge pressures, unless otherwise indicated.
Preferably, the drying gas is obtained by treatment in a dehydration tower, and the water content in the drying gas is 0.1-100 ppm.
Preferably, an outlet heat exchanger I is connected outside the carbonylation reactor; industrial grade CO, N2And the drying gas from the dehydration column as carbonylThe carbonylation reaction raw material exchanges heat with the carbonylation reaction product from the carbonylation reactor through the outlet heat exchanger I and then enters the carbonylation reactor for carbonylation reaction.
Preferably, part of the gas phase component from the top of the methanol washing tower is pressurized by a carbonylation circulating compressor and then enters the esterification reaction tower.
Preferably, the gas-phase light component at the top of the NO recovery tower enters the MN recovery tower after being compressed and pressurized by a compressor.
Preferably, an outlet heat exchanger II is connected outside the hydrogenation reactor; and the dimethyl oxalate component from the DMO rectifying tower, the industrial hydrogen and the recycle gas from the pressurized recycle compressor and the recycle gas from the membrane separator are used as hydrogenation reaction raw materials, and the raw materials and the hydrogenation products from the hydrogenation reactor exchange heat through the outlet heat exchanger II and then enter the hydrogenation reactor for hydrogenation reaction.
Preferably, the liquid phase separated by the second gas-liquid separator firstly enters a methanol separation tower; the non-condensable gas recovered from the top of the methanol separation tower enters the membrane separator, part of liquid phase light components such as methanol recovered from the top of the methanol separation tower enter the upper part of the methanol washing tower to be used as washing liquid, and part of the liquid phase light components enter the NO recovery tower; the heavy component of the liquid phase at the bottom of the methanol separation tower enters a light component rectifying tower for further separation and purification; the light component at the top of the light component rectifying tower enters an out-of-range alcohol recovery device for recovery treatment; heavy components in the tower bottom of the light component rectifying tower enter the ethylene glycol product tower; and (3) the light components at the top of the ethylene glycol product tower enter an out-of-the-home 1, 2-BDO recovery processing device for further recovery processing, the heavy components at the bottom of the ethylene glycol product tower enter the out-of-the-home recovery processing device for subsequent processing, and the ethylene glycol product is led out from the upper side line of the ethylene glycol product tower.
Preferably, the second gas-liquid separator comprises a high-pressure gas-liquid separator and a low-pressure gas-liquid separator; the gas phase separated by the high-pressure gas-liquid separator enters the hydrogenation circulating compressor, and the gas phase enters the low-pressure gas-liquid separator; the liquid phase separated by the high-pressure gas-liquid separator enters the methanol separation tower; and the gas phase separated by the low-pressure gas-liquid separator enters the membrane separator, and the liquid phase separated by the low-pressure gas-liquid separator enters the methanol separation tower.
Preferably, 0.1-10 v% of the gas phase separated by the high-pressure gas-liquid separator enters the low-pressure gas-liquid separator.
Preferably, the gas phase separated by the low-pressure gas-liquid separator and the non-condensable gas from the top of the methanol separation tower enter the membrane separator after methanol is absorbed by the methanol absorption tank.
Preferably, the carbonylation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor.
More preferably, the carbonylation plate reactor is a plate fixed bed carbonylation reactor.
Preferably, a plate group fixing cavity is arranged in the center of the plate type fixed bed carbonylation reactor, and a plate group is arranged in the plate group fixing cavity; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed carbonylation reactor; the catalyst bed layer is filled with a carbonylation reaction catalyst; after the carbonylation reaction raw material reaches the inlet temperature of the catalyst bed layer, the carbonylation reaction raw material enters the catalyst bed layer from the top of the plate type fixed bed carbonylation reactor to carry out carbonylation reaction; the method comprises the following steps that a refrigerant introduced from the outside enters a plate group fixing cavity from the bottom of a plate type fixed bed carbonylation reactor and is led out from the top of the plate type fixed bed carbonylation reactor, and the heat exchange is carried out in the countercurrent process to take away the reaction heat of the carbonylation reaction; the carbonylation product from the bottom of the catalyst bed is withdrawn from the bottom of the plate fixed bed carbonylation reactor.
Preferably, a steam drum I is connected outside the plate type fixed bed carbonylation reactor; the method comprises the following steps that (1) a refrigerant introduced from the outside enters a steam pocket I, and the refrigerant in the steam pocket I enters a plate set fixing cavity of a plate type fixed bed carbonylation reactor to exchange heat with a catalyst bed layer, so that reaction heat is removed; the heated refrigerant is a vapor-liquid mixture, enters the steam drum I for gas-liquid separation, and the generated low-pressure saturated steam enters the outdoor low-pressure steam recovery system for recycling.
Preferably, the carbonylation catalyst is a catalyst sold in Shanghai Peng Zheng engineering technology, and the trade name of the catalyst is DMO-0701T.
Preferably, the esterification reaction tower is a packed tower;
preferably, the esterification reaction column is a tray-filler mixing column having both a tray portion and a filler-packed portion.
Preferably, the number of theoretical plates of the esterification reaction tower is 20-50. The tower plates are sequentially arranged from the top to the bottom in sequence.
Preferably, in the feed of the esterification reaction tower, the O2Feeding from 16 th to 50 th tower plates respectively in 2-8 paths; feeding the NO and the overhead gas phase light components from the methanol washing tower from a tower plate 18-50; feeding the fresh methanol, the recovered methanol from the top of the methanol recovery tower, the mixture of the methanol and the methyl nitrite recovered from the top of the methanol rectification tower and the methyl nitrite-containing alcoholic solution from the tower kettle of the MN recovery tower from the 1 st to 5 th tower plates; feeding the reflux material of the tower kettle of the esterification reaction tower from a 10 th to a 25 th tower plate.
Preferably, in the esterification reaction tower, O2The molar ratio of NO to methanol is 0.01-0.8: 0.1-3.2: 0.8 to 50.
Preferably, the temperature of the top of the esterification reaction tower is 30-80 ℃, the temperature of the tower kettle is 50-200 ℃, the temperature of the reaction zone is 50-160 ℃, and the reaction pressure is 0.5-10 MPa.
Preferably, the methanol recovery tower, the methanol washing tower, the methanol rectifying tower, the nitric acid concentrating tower, the NO recovery tower, the MN recovery tower and the DMO rectifying tower are packed towers, plate towers or bubble cap towers.
Preferably, the filler filled in the packed tower is random packing or efficient regular packing; the random packing is in a saddle shape, a Raschig ring, a pall ring, a wheel shape, a rectangular saddle ring, a spherical shape or a columnar shape; the efficient structured packing is corrugated packing, grid packing or pulse packing.
Preferably, the theoretical plate number of the methanol recovery tower is 5-50, the tower top temperature is 40-150 ℃, the tower kettle temperature is 60-230 ℃, and the tower top pressure is 0.01-2.0 MPa.
Preferably, the reflux ratio of the light components at the top of the methanol recovery tower is 0.1-3.0.
Preferably, the proportion of the part which is recycled into the esterification reaction tower in the methanol recovery tower top to recover the methanol is 10-90 wt%.
Preferably, the theoretical plate number of the methanol washing tower is 10-50, the tower top temperature is 15-70 ℃, the tower kettle temperature is 10-100 ℃, and the tower top pressure is 0.9-10 MPa.
Preferably, in the gas phase component at the top of the methanol washing tower, the proportion of the purge gas is 0.05-5 v%.
Preferably, the methanol rectifying tower is an extraction rectifying tower, the number of theoretical plates is 10-60, the temperature of the top of the tower is 50-150 ℃, the temperature of a tower kettle is 130-250 ℃, and the pressure of the top of the tower is 0.01-0.5 MPa.
Preferably, the theoretical plate number of the nitric acid concentration tower is 1-30, the tower top temperature is 30-110 ℃, the tower kettle temperature is 60-160 ℃, and the tower top pressure is 0.01-0.3 MPa.
Preferably, the reflux ratio of the light components at the top of the nitric acid concentration tower is 0.01-3.
Preferably, the number of theoretical plates of the NO recovery tower is 5-30, the temperature of the top of the tower is 30-120 ℃, the temperature of a tower kettle is 50-200 ℃, and the pressure of the top of the tower is 1-10 MPa.
Preferably, the purge gas is fed from 5 th to 30 th tower plates of the NO recovery tower; feeding the concentrated nitric acid from a tower plate 1-10 of an NO recovery tower; the recovered methanol from the top of the methanol separation tower is fed from the 1 st to 10 th tower plates.
Preferably, in the NO recovery tower, the molar ratio of the nitric acid to the methanol to the NO in the purge gas is 1.1-10: 2-100: 1-5.
Preferably, the number of theoretical plates of the MN recovery tower is 10-60, the temperature of the top of the tower is 0-50 ℃, the temperature of the bottom of the tower is 0-80 ℃, and the reaction pressure is 1-10 MPa.
Preferably, the DMO rectifying tower has 10-50 theoretical plates, the tower top temperature is 80-120 ℃, the tower kettle temperature is 120-200 ℃, and the operation is carried out under normal pressure or reduced pressure.
Preferably, the reflux ratio of light components at the top of the DMO rectifying tower is 0.1-100.
Preferably, the composition of the purge gas recovered in the pressure swing adsorption tank is: n is a radical of260-80 v% of CO, and 20-40 v% of CO; separated CO2The gas accounts for 0.1-5 v% of the total amount of the intake air, wherein CO2The concentration of (A) is 99.8-99.9 v%; separated CO2The gas may be processed by an off-site device.
Preferably, the hydrogenation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor.
More preferably, the hydrogenation plate type reactor is a plate type fixed bed hydrogenation reactor.
Preferably, a plate group fixing cavity is arranged in the center of the plate type fixed bed hydrogenation reactor, and a plate group is arranged in the plate group fixing cavity; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed hydrogenation reactor; the catalyst bed layer is filled with a hydrogenation reaction catalyst; after the hydrogenation reaction raw material reaches the inlet temperature of the catalyst bed, the hydrogenation reaction raw material enters the catalyst bed from the top of the plate-type fixed bed hydrogenation reactor to carry out hydrogenation reaction; a refrigerant introduced from the outside enters the plate group fixing cavity from the bottom of the plate type fixed bed hydrogenation reactor and is led out from the top of the plate type fixed bed hydrogenation reactor, and the heat exchange is carried out in the countercurrent process to take away the reaction heat of the hydrogenation reaction; the hydrogenation product from the bottom of the catalyst bed is drawn from the bottom of the plate fixed bed hydrogenation reactor.
Preferably, a steam drum II is connected outside the plate-type fixed bed hydrogenation reactor; the method comprises the following steps that (1) a refrigerant introduced from the outside enters a steam pocket II, the refrigerant in the steam pocket II enters a plate group fixing cavity of a plate type fixed bed hydrogenation reactor to exchange heat with a catalyst bed layer, and reaction heat is removed; the heated refrigerant is a vapor-liquid mixture, enters a steam drum II for gas-liquid separation, and the generated low-pressure saturated steam enters an outdoor low-pressure steam recovery system for recycling.
Preferably, the refrigerant is water or heat transfer oil, and is preferably water.
Preferably, a start-up heater is connected outside the plate-type fixed bed hydrogenation reactor; in the initial startup stage, the temperature cannot meet the reaction requirement, the hydrogenation reaction raw material enters a startup heater for preheating, and enters a catalyst bed layer for hydrogenation reaction after the preheating reaches the inlet temperature of the catalyst bed layer; in the initial startup stage, the startup heater provides a unique heat source for the hydrogenation reaction in the plate-type fixed bed hydrogenation reactor; the heat source of the start-up heater is low-pressure steam.
Preferably, the hydrogenation catalyst is selected from catalysts sold in Shanghai Peng engineering technology, Inc., and the catalyst is sold under the trade name MEG-801T.
Preferably, the theoretical plate number of the methanol separation tower is 10-40, the tower top temperature is 40-70 ℃, the tower kettle temperature is 80-180 ℃, and the operation is carried out under normal pressure or reduced pressure; the reflux ratio of light components at the top of the methanol separation tower is 0.1-3.
Preferably, the number of theoretical plates of the light component rectifying tower is 10-60, the temperature of the top of the tower is 58-90 ℃, the temperature of a tower kettle is 70-160 ℃, and the absolute pressure of the top of the tower is 5-50 KPa.
Preferably, the reflux ratio of the light components at the top of the light component rectifying tower is 1-50.
Preferably, the number of theoretical plates of the ethylene glycol product tower is 30-100, the temperature of the top of the tower is 100-150 ℃, the temperature of a tower kettle is 130-230 ℃, and the absolute pressure of the top of the tower is 5-50 KPa; the reflux ratio of light components at the top of the ethylene glycol product tower is 50-120 or total reflux.
Preferably, the membrane separator is formed by connecting 1-100 hollow fiber membrane modules in parallel or in series.
Preferably, the tolerant pressure of the tube shell of the membrane separator is 4.75MPa, and the maximum pressure difference is 1.5MPa (raw material gas to permeation gas); the operating temperature of the membrane separator was 85 ℃ at the most.
Preferably, the concentration of hydrogen in the purified gas obtained by separation and purification through a membrane separator is 88-99.00 v%, and the recovery rate of hydrogen is 90-98.5%.
The basic principle of the membrane separator is that the partial pressure difference of gases at two sides of a hollow fiber membrane is used as a driving force, and the purpose of separation is achieved by utilizing different selective permeabilities of the hollow fiber membrane to various gases through the steps of permeation, dissolution, diffusion, analysis and the like. The raw material gas flows through the shell side of the hollow fiber membrane component, the permeating gas flows through the tube side, and the tail gas enters the next hollow fiber membrane component. Due to H2The permeation rate at the membrane surface is CH4、N2Ar, etc. are several tens of times so H2Enters each hollow fiber pipe, is collected and then is discharged from the lower part of the membrane separator, and non-permeable gas (tail gas) is discharged from the upper part of the hollow fiber membrane component. The hollow fiber membrane component is internally provided with a core member consisting of 1000-100000 hollow fiber membrane filament tubes, and the fiber tubes are formed by specially processing high polymer materials. The raw material gas enters from the side port of the separator, and when the gas flowing downwards along the outer side of the fiber tube bundle contacts with the outer surface of the fiber membrane silk tube, the gas is dissolved, permeated and diffused on the fiber wall, and different gases are separated by utilizing the difference of the dissolving and permeating capacities of various gases.
The invention has the technical effects and advantages that:
because the carbonylation system and the esterification system adopt high-pressure operation, the requirement of large-scale synthesis gas to ethylene glycol process devices on the volume of equipment can be greatly reduced, the large-scale production of a single series of devices is facilitated, the safe production of the devices is facilitated, and the equipment investment is reduced.
The nitric acid waste liquid recycling process and the purge gas recycling process are highly coupled, and waste liquid generated in the device can be recycled as a raw material for recovering a large amount of nitric oxide purge gas to generate methyl nitrite required by the main reaction. The process combination technology is scientific and reasonable, the full recycling of the discharged waste gas and waste liquid is realized through one reactor, and the process is economical and environment-friendly.
The carbonylation plate type reactor is a plate type reactor, realizes the reaction of preparing the dimethyl oxalate by CO carbonylation coupling, can fully utilize the characteristic of uniform temperature distribution of the reactor, and achieves the characteristics of improving the space-time yield of the dimethyl oxalate and recycling reaction heat. Meanwhile, the utilization coefficient of the catalyst and the utilization rate of the volume of the reactor are improved, the filling amount of the catalyst is increased, and the production capacity of the reactor is improved. The reaction characteristics also obtain the same effects of energy conservation and consumption reduction in the preparation of the ethylene glycol by the hydrogenation of the dimethyl oxalate.
The recovery of the purge gas in the hydrogenation section fully saves precious hydrogen resources, further reduces unit coal consumption, is beneficial to reducing the overall energy consumption and pollution emission of the device, and has realistic significance. Meanwhile, the recovery of the purge gas of the hydrogenation section in the process can reduce the pressure of a reaction system by about 1MPa under the same load by adopting a membrane separation system, and for a compression system, the reduction of the outlet pressure can save a large amount of power consumption. The valuable hydrogen resources are fully saved, the unit coal consumption is reduced, the overall energy consumption and the pollution emission of the device are reduced, and the method has realistic significance. By adopting the membrane separation system, the hydrogenation reaction rate is favorably improved, and the daily yield of the ethylene glycol is increased by about 10 percent compared with the original daily yield.
In conclusion, by adopting the high-pressure process flow and the plate type reactor, the bottleneck of large-scale device is effectively solved, the equipment investment is reduced, the waste heat of reaction heat is recovered, the energy consumption of unit ethylene glycol production is reduced, and the consumption of steam and cooling water is reduced; through the process coupling of waste gas and waste liquid, the emission of toxic substances is reduced, thereby achieving the dual purposes of energy conservation and environmental protection. The invention realizes the full recycling of the discharged waste gas and waste liquid, the comprehensive energy application of the reaction heat of the device and the tower separation, improves the energy utilization efficiency, saves the energy consumption and has obvious industrial application value. The invention provides guarantee for the technical development of preparing the ethylene glycol from the synthesis gas to the more environment-friendly, more efficient and more energy-saving technology. The invention is technically feasible and economically reasonable.
The process optimization design can obviously improve the yield, and is not described in any literature. The process provided by the invention is particularly favorable from the energy consumption perspective, has the characteristic of remarkably saving energy consumption, and has very remarkable effects by combining the application of the useful substance recycling step, particularly the high coupling of the nitric acid waste liquid recycling process and the purge gas recycling process, the separation process of the nitric acid waste liquid recycling process and the recycling of hydrogen in reaction waste gas.
Drawings
FIG. 1 shows a system (part) of an apparatus for producing dimethyl oxalate and hydrogenating ethylene glycol by high pressure carbonylation of industrial synthesis gas
FIG. 2 shows a system (part) of an apparatus for producing dimethyl oxalate by high pressure carbonylation of industrial synthesis gas and producing ethylene glycol by hydrogenation
Reference numerals:
1, a carbonylation reactor; 2, a steam drum I; 3; an outlet heat exchanger I; 4, a first gas-liquid separator; 5, a methanol rectifying tower; 6, DMO rectifying tower; 7, a methanol washing tower; 8, a carbonylation recycle compressor; 9, an esterification reaction tower; 10, a dehydration tower; 11, a methanol recovery tower; 12, a nitric acid concentration tower; 13, an NO recovery column; 14, a compressor; 15, MN recovery column; 16, a pressure swing adsorption tank; 17, a hydrogenation reactor; 18, a steam drum II; 19, starting a heater; 20, an outlet heat exchanger II; 21, high-pressure gas-liquid separator; 22, a methanol separation tower; 23, a light component rectifying tower; 24, an ethylene glycol product column; 25, a hydrogenation recycle compressor; 26, a low pressure gas-liquid separator; 27, a methanol absorption tank; 28, a membrane separator.
Detailed Description
The technical solution of the present invention is illustrated by specific examples below. It is to be understood that one or more method steps mentioned in the present invention do not exclude the presence of other method steps before or after the combination step or that other method steps may be inserted between the explicitly mentioned steps; it should also be understood that these examples are for illustrative purposes only and are not intended to limit the scope of the present invention. Moreover, unless otherwise indicated, the numbering of the various method steps is merely a convenient tool for identifying the various method steps, and is not intended to limit the order in which the method steps are arranged or the scope of the invention in which the invention may be practiced, and changes or modifications in the relative relationship may be made without substantially changing the technical content.
Experimental procedures without specific conditions noted in the examples below, generally following conventional conditions, such as: a manual for chemical operations, or according to the conditions recommended by the manufacturer.
As shown in fig. 1 and fig. 2, a device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation comprises a carbonylation reaction system, an esterification reaction system, a purge gas and waste acid coupling recovery system and a hydrogenation reaction system;
the carbonylation reaction system comprises a carbonylation reactor 1, a first gas-liquid separator 4, a methanol washing tower 7, a methanol rectifying tower 5 and a DMO rectifying tower 6; the carbonylation reactor 1 is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the first gas-liquid separator 4 is provided with a feed inlet, a gas outlet and a liquid outlet; the methanol washing tower 7 is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the methanol rectifying tower 5 is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the DMO rectifying tower 6 is provided with a lower feed inlet, a top outlet and a bottom outlet;
the esterification reaction system comprises an esterification reaction tower 9 and a methanol recovery tower 11; the esterification reaction tower 9 is provided with a top feed inlet, an upper feed inlet, a plurality of lower feed inlets, a middle reflux inlet, a top outlet and a bottom outlet; the methanol recovery tower 11 is provided with a middle lower part feed inlet, a top outlet and a bottom outlet;
the purge gas and waste acid coupling recovery system comprises a nitric acid concentration tower 12, an NO recovery tower 13, an MN recovery tower 15 and a pressure swing adsorption tank 16; the nitric acid concentration tower 12 is provided with a middle feeding hole, a top outlet and a bottom outlet; the NO recovery tower 13 is provided with a top feed inlet, a middle feed inlet, a bottom feed inlet, a top outlet and a bottom outlet; the MN recovery tower 15 is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the pressure swing adsorption tank 16 is provided with a feed inlet, a recovered gas outlet and an exhaust gas outlet;
the hydrogenation reaction system comprises a hydrogenation circulating compressor 14, a hydrogenation reactor 17, a second gas-liquid separator, a membrane separator 28, a methanol separation tower 22, a light component rectifying tower 23 and an ethylene glycol product tower 24; the hydrocycle compressor 14 comprises an inlet and an outlet; the hydrogenation plate reactor 17 is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the second gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the membrane separator 28 is provided with a feed inlet, a recovered gas outlet and a discharged gas outlet; the methanol separation tower 22 is provided with a middle feed inlet, a top noncondensable gas outlet, a top liquid phase light component outlet and a bottom liquid phase heavy component outlet; the light component rectifying tower 23 is provided with a lower feed inlet, a top outlet and a bottom outlet; the ethylene glycol product tower 24 is provided with a lower feed inlet, a top outlet, an upper outlet and a bottom outlet;
the top feed inlet of the carbonylation reactor 1 is connected with a CO raw material pipeline and N2The raw material pipelines are connected through pipelines; the bottom discharge hole of the carbonylation reactor 1 is connected with the feed inlet of the first gas-liquid separator 4 through a pipeline; a gas outlet of the first gas-liquid separator 4 is connected with a lower feed inlet of the methanol washing tower 7 through a pipeline; a liquid outlet of the first gas-liquid separator 4 is connected with an upper feed inlet of the methanol rectifying tower 5 through a pipeline; a branch outlet A and a branch outlet B are arranged at the outlet at the top of the methanol washing tower 7, the branch outlet A is connected with a lower feed inlet of the esterification reaction tower 9 through a pipeline, and the branch outlet B is connected with a bottom feed inlet of the NO recovery tower 13 through a pipeline; the bottom outlet of the methanol washing tower 7 is connected with the lower feed inlet of the methanol rectifying tower 5 through a pipeline; the top outlet of the methanol rectifying tower 5 is connected with the upper feed inlet of the esterification reaction tower 9 through a pipeline; the bottom outlet of the methanol rectifying tower 5 is connected with the lower feed inlet of the DMO rectifying tower 6 through a pipeline; the bottom outlet of the DMO rectifying tower 6 is connected with the top feed inlet of the hydrogenation reactor 17 through a pipeline, and the top outlet of the DMO rectifying tower 6 is a DMC product outlet;
other lower feed inlets of the esterification reaction tower 9, an NO raw material pipeline and a multi-path O2The raw material pipelines are respectively connected through pipelines; a feed inlet at the top of the esterification reaction tower 9 is connected with a methanol raw material pipeline through a pipeline; a branch outlet C and a branch outlet D are arranged at the bottom outlet of the esterification reaction tower 9, the branch outlet C is connected with the middle reflux inlet of the esterification reaction tower 9 through a pipeline, and the branch outlet D is connected with the lower feeding port of the methanol recovery tower 11 through a pipeline; the top outlet of the esterification reaction tower 9 is connected with the top feed inlet of the carbonylation reactor 1 through a pipeline; the top outlet of the methanol recovery tower 11 is provided with a branch outlet E and a branch outlet FA branch outlet E is connected with an upper feed inlet of the esterification reaction tower 9 through a pipeline, and a branch outlet F is connected with an upper feed inlet of the MN recovery tower 15 through a pipeline; the bottom outlet of the methanol recovery tower 11 is connected with the middle feed inlet of the nitric acid concentration tower 12 through a pipeline;
the top outlet of the nitric acid concentration tower 12 is a waste liquid outlet; the bottom outlet of the nitric acid concentration tower 12 is connected with the middle feed inlet of the NO recovery tower 13 through a pipeline; the top outlet of the NO recovery tower 13 is connected with the lower feed inlet of the MN recovery tower 15 through a pipeline; the bottom outlet of the NO recovery tower 13 is connected with the feeding port at the middle lower part of the methanol recovery tower 11 through a pipeline; the top outlet of the MN recovery tower 15 is connected with the feed inlet of the pressure swing adsorption tank 16 through a pipeline; the bottom outlet of the MN recovery tower 15 is connected with the upper feed inlet of the esterification reaction tower 9 through a pipeline; the recycled gas outlet of the pressure swing adsorption tank 16 is connected with the top feed inlet of the carbonylation plate reactor 1 through a pipeline; the exhaust gas outlet of the pressure swing adsorption tank 16 is connected with an external recovery device through a pipeline;
the inlet of the hydrogenation circulation compressor 14 is connected with an industrial hydrogen raw material pipeline through a pipeline, and the outlet of the hydrogenation circulation compressor 14 is connected with the top feed inlet of the hydrogenation reactor 17 through a pipeline; a discharge hole at the bottom of the hydrogenation reactor 17 is connected with a feed hole of the second gas-liquid separator through a pipeline; a gas outlet of the second gas-liquid separator is provided with a branch outlet G and a branch outlet H, the branch outlet G is connected with an inlet of the hydrogenation circulating compressor 14 through a pipeline, and the branch outlet H is connected with a feed inlet of the membrane separator 28 through a pipeline; a liquid outlet of the second gas-liquid separator is connected with a lower feed port of the methanol separation column 22 through a pipeline; the top noncondensable gas outlet of the methanol separation tower 22 is connected with the feed inlet of the membrane separator 28 through a pipeline; a branch outlet I and a branch outlet J are arranged at a top liquid phase light component outlet of the methanol separation tower 22, the branch outlet I is connected with an upper feed inlet of the methanol washing tower 7 through a pipeline, and the branch outlet J is connected with a top feed inlet of the NO recovery tower 13 through a pipeline; a bottom liquid phase heavy component outlet of the methanol separation tower 22 is connected with a lower feed inlet of the light component rectifying tower 23 through a pipeline; the top light component outlet of the light component rectifying tower 23 is connected with an out-of-range alcohol recovery device through a pipeline; the bottom heavy component outlet of the light component rectifying tower 23 is connected with the lower feed inlet of the ethylene glycol product tower 24 through a pipeline; the top outlet of the ethylene glycol product tower 24 is connected with an extra-terrestrial 1, 2-BDO recovery and treatment device through a pipeline; the bottom outlet of the ethylene glycol product tower 24 is connected with an outside recovery processing device through a pipeline; the upper outlet of the ethylene glycol product tower 24 is an ethylene glycol product outlet; the exhaust gas outlet of the membrane separator 28 is connected with an out-of-range recovery device through a pipeline, and the recovered gas outlet of the membrane separator 28 is connected with the top feed inlet of the hydrogenation reactor 17 through a pipeline.
In a preferred embodiment, the carbonylation reactor 1 is externally connected with a dehydration column 10; the dehydration tower 10 is provided with a feed inlet and a dry gas outlet; the top outlet of the esterification reaction tower 9 and the recovered gas outlet of the pressure swing adsorption tank 16 are connected with the feed inlet of the dehydration tower 10 through pipelines; the dry gas outlet of the dehydration tower 10 is connected with the feed inlet at the top of the carbonylation reactor 1 through a pipeline.
The dehydration tower consists of two molecular sieve dryers A and B which alternately run and regenerate; the molecular sieve dryer A and the molecular sieve dryer B are filled with adsorbents.
In a preferred embodiment, the bottom discharge port of the carbonylation reactor 1 is connected with an outlet heat exchanger I3; the outlet heat exchanger I3 is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the CO raw material pipeline and N2A raw material pipeline and a dry gas outlet of the dehydration tower 10 are connected with a cold material flow inlet of the outlet heat exchanger I3 through pipelines; a cold material flow outlet of the outlet heat exchanger I3 is connected with a top feed inlet of the carbonylation reactor 1 through a pipeline; a discharge port at the bottom of the carbonylation reactor 1 is connected with a hot material flow inlet of the outlet heat exchanger I3 through a pipeline; the hot material outlet of the outlet heat exchanger I3 is separated from the first gas-liquidThe inlet ports of the vessel 4 are connected via a pipeline.
In a preferred embodiment, the carbonylation reactor 1 is externally connected with a steam drum I2; the steam pocket I2 is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a vapor outlet; the refrigerant inlet of the steam pocket I2 is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum I2 is connected with a refrigerant inlet at the bottom of the carbonylation plate type reactor 1 through a pipeline; a top refrigerant outlet of the carbonylation reactor 1 is connected with a vapor-liquid mixture inlet of the steam drum I2 through a pipeline; and a steam outlet of the steam drum I2 is connected with an out-of-range steam recovery system through a pipeline.
As a preferred embodiment, a carbonylation circulating compressor 8 is connected between the branched outlet a of the methanol washing tower 7 and the lower feed inlet of the esterification reaction tower 9; the carbonylation circulating compressor 8 is provided with an inlet and an outlet; the branch outlet A is connected with the inlet of the carbonylation circulating compressor 8 through a pipeline; the outlet of the carbonylation circulating compressor 8 is connected with the lower feeding port of the esterification reaction tower 9 through a pipeline.
In a preferred embodiment, a compressor 14 is connected between the top outlet of the NO recovery tower 13 and the bottom feed inlet of the MN recovery tower 15; the compressor 14 is provided with an inlet and an outlet; the top outlet of the NO recovery tower 13 is connected with the inlet of the compressor 14 through a pipeline; the outlet of the compressor is connected with the bottom feed inlet of the MN recovery tower 15 through a pipeline.
As a preferred embodiment, the bottom discharge hole of the hydrogenation reactor 17 is connected with an outlet heat exchanger ii 20; the outlet heat exchanger II 20 is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the bottom outlet of the DMO rectifying tower 6, the recovered gas outlet of the membrane separator 28 and the outlet of the hydrogenation circulating compressor 25 are connected with the cold material flow inlet of the outlet heat exchanger II 20 through pipelines; a cold material flow outlet of the outlet heat exchanger II 20 is connected with a top feeding hole of the hydrogenation reactor 17 through a pipeline; a discharge hole at the bottom of the hydrogenation reactor 17 is connected with a hot material flow inlet of the outlet heat exchanger II 20 through a pipeline; and the hot material flow outlet of the outlet heat exchanger II 20 is connected with the feed inlet of the second gas-liquid separator through a pipeline.
In a preferred embodiment, a start-up heater 19 is connected to the top feed inlet of the hydrogenation reactor 17; the start-up heater 19 is provided with a feed inlet and a discharge outlet; a cold material flow outlet of the outlet heat exchanger II 20 is connected with a feed inlet of the start-up heater 19 through a pipeline; the discharge hole of the start-up heater is connected with the top feed inlet of the hydrogenation reactor 17 through a pipeline.
In a preferred embodiment, a steam drum II 18 is connected outside the hydrogenation reactor 17; the steam drum II 18 is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a vapor outlet; the refrigerant inlet of the steam drum II 18 is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum II 18 is connected with a refrigerant inlet at the bottom of the hydrogenation reactor 17 through a pipeline; a top refrigerant outlet of the hydrogenation reactor 17 is connected with a vapor-liquid mixture inlet of the steam drum II 18 through a pipeline; and a steam outlet of the steam drum II 18 is connected with an out-of-range steam recovery system through a pipeline.
As a preferred embodiment, the second gas-liquid separator includes a high-pressure gas-liquid separator 21 and a low-pressure gas-liquid separator 26; the high-pressure gas-liquid separator 21 is provided with a feed inlet, a gas outlet and a liquid outlet; the low-pressure gas-liquid separator 26 is provided with a feed inlet, a gas outlet and a liquid outlet; a discharge hole at the bottom of the hydrogenation reactor 17 is connected with a feed hole of the high-pressure gas-liquid separator 21 through a pipeline; a gas outlet of the high-pressure gas-liquid separator 21 is provided with a branch outlet K and a branch outlet L, the branch outlet K is connected with an inlet of the hydrogenation circulating compressor 25 through a pipeline, and the branch outlet L is connected with a feed inlet of the low-pressure gas-liquid separator 26 through a pipeline; the liquid outlet of the high-pressure gas-liquid separator 21 is connected with the middle feed inlet of the methanol separation tower 22 through a pipeline; the gas outlet of the low-pressure gas-liquid separator 26 is connected with the feed inlet of the membrane separator 28 through a pipeline; the liquid outlet of the low-pressure gas-liquid separator 26 is connected with the middle feed inlet of the methanol separation tower 22 through a pipeline.
In a preferred embodiment, a methanol absorption tank 27 is arranged in front of the feed inlet of the membrane separator 28; the methanol absorption tank 27 is provided with a feed inlet and a purified gas outlet; a gas outlet of the low-pressure gas-liquid separator 26 and a top noncondensable gas outlet of the methanol separation tower 22 are connected with a feed inlet of the methanol absorption tank 27 through pipelines; a purified gas outlet of the methanol absorption tank 27 is connected to a feed port of the membrane separator 28 via a line.
The carbonylation reactor 1 can be a plate reactor, a tubular reactor or a tubular-plate composite reactor;
as a preferred embodiment, the carbonylation reactor 1 is a plate-type fixed bed carbonylation reactor;
the center of the plate type fixed bed carbonylation reactor is provided with a plate group fixing cavity, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed carbonylation reactor; the catalyst bed layer is filled with a carbonylation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate type fixed bed carbonylation reactor, a bottom refrigerant inlet of the plate type fixed bed carbonylation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate type fixed bed carbonylation reactor through a pipeline; and the top inlet of the plate fixed bed carbonylation reactor is connected with the top inlet of the catalyst bed layer through a pipeline, and the top outlet of the plate group fixing cavity is connected with the top refrigerant outlet of the plate fixed bed carbonylation reactor through a pipeline.
As a preferred embodiment, the esterification reaction tower 9 is a packed tower;
as a more preferred embodiment, the esterification reaction column 9 is a tray-filler mixing column having both a tray portion and a filler-filled portion.
In a preferred embodiment, the methanol washing column 7, the methanol rectifying column 5, the methanol recovering column 11, the NO recovering column 13, the MN recovering column 15, the DMO rectifying column 6, and the nitric acid concentrating column 12 are packed columns, plate columns, or bubble column.
As a preferred embodiment, the packing filled in the packed tower is random packing or high-efficiency structured packing; the random packing is in a saddle shape, a Raschig ring, a pall ring, a wheel shape, a rectangular saddle ring, a spherical shape or a columnar shape; the efficient regular packing is corrugated packing, grid packing and pulse packing. .
The hydrogenation plate type reactor 17 can be a plate type reactor, a tubular reactor or a tubular-plate type composite reactor;
as a preferred embodiment, the hydrogenation reactor 17 is a plate-type fixed bed hydrogenation reactor;
a plate group fixing cavity is arranged in the center of the plate type fixed bed hydrogenation reactor, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed hydrogenation reactor; the catalyst bed layer is filled with a hydrogenation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate-type fixed bed hydrogenation reactor, a bottom refrigerant inlet of the plate-type fixed bed hydrogenation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate-type fixed bed hydrogenation reactor through a pipeline; the top of plate fixed bed hydrogenation ware, plate fixed bed hydrogenation ware's top feed inlet with the top entry pipe line connection of catalyst bed layer, the top export in the fixed chamber of plate group with plate fixed bed hydrogenation ware's top refrigerant export pipe line connection.
As a preferred embodiment, the membrane separator 28 is composed of 1-100 hollow fiber membrane modules connected in parallel or in series.
As shown in fig. 1 and fig. 2, the process flow for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation provided by the invention is as follows:
NO from line 18, fresh methanol from line 26, and O in 2-8 feeds2Performing gas-liquid countercurrent contact in esterification tower 9 to perform esterification reaction, collecting MN mixed gas generated at the top of the tower with the recovered gas phase from the pressure swing adsorption tank of pipeline 39 via pipeline 23, introducing into dehydration tower 10 via pipeline 24 for dehydration, and passing the dehydrated dry gas through pipeline 25, CO from pipeline 1 and N from pipeline 22Mixed and then enters a pipeline 3 as the raw material gas for the carbonylation reaction. The acid waste liquid containing a large amount of methanol flows back to the esterification reaction tower 9 through a pipeline 20 in a certain amount in a tower kettle in the esterification reaction tower 9, and the rest acid waste liquid enters a methanol recovery tower 11 through a pipeline 21 and the methanol acid waste liquid from a pipeline 33 at the same time for methanol recovery; the light components of methanol generated at the top of the methanol recovery tower 11 are shunted after passing through a pipeline 28, except that a part of the light components enter the MN recovery tower 15 through a pipeline 29 to be used as washing liquid, the other part of the light components is converged with fresh methanol from a pipeline 26 and is used as an alcohol source of the esterification reaction tower 9 through a pipeline 22; acid-containing wastewater generated in the tower bottom of the methanol recovery tower 11 enters a nitric acid concentration tower 12 through a pipeline 27 for nitric acid concentration.
The carbonylation reaction raw material from the pipeline 3 exchanges heat with the carbonylation reaction product discharged from the bottom of the carbonylation reactor 1 through an outlet heat exchanger I3, and then enters a catalyst bed layer from the top of the carbonylation reactor 1 for carbonylation reaction; meanwhile, refined water from the outside of the system enters the steam pocket I2 through a pipeline 8, a refrigerant in the steam pocket I2 enters the plate group fixing cavity from the bottom of the carbonylation reactor 1 through a pipeline 9 to exchange heat with a catalyst bed layer, heat generated by reaction is removed, the heated refrigerant is a vapor-liquid mixture, the vapor-liquid mixture is led out from the top of the carbonylation reactor 1 and then enters the steam pocket I2 to be subjected to gas-liquid separation, and generated low-pressure saturated steam enters an outdoor low-pressure steam recovery system through a pipeline 7 to be recycled. The carbonylation reaction product enters a first gas-liquid separator 4 for gas-liquid separation after heat exchange through an outlet heat exchanger I3, and a gas phase component containing most DMC (dimethyl carbonate) enters a methanol washing tower 7 through a pipeline 11 and is in countercurrent contact with recovered methanol from a pipeline 57; the DMO heavy component at the tower bottom of the first gas-liquid separator 4 enters a methanol rectifying tower 5 through a pipeline 10 and methanol washing liquid containing MN (methyl nitrite), DMC and DMO (dimethyl oxalate) in the tower bottom of a methanol washing tower 7 through a pipeline 12, and the two streams are in countercurrent contact for extraction and separation; most of gas phase light components at the top of the methanol washing tower 7 enter the esterification reaction tower 9 through a carbonylation circulating compressor 8 through a pipeline 17 for recycling, and a small part of gas phase light components are used as purge gas and enter the NO recovery tower 13 through a pipeline 32 for recovery treatment; the mixture of methanol and methyl nitrite recovered from the top of the methanol rectifying tower 5 is recycled to the esterification reaction tower 9 through a pipeline 14 for reuse, and heavy components in the tower kettle enter the DMO rectifying tower 6 through a pipeline 13; DMC products are obtained at the top of the DMO rectifying tower 6, and dimethyl oxalate components at the bottom of the tower enter a pipeline 15 to be used as raw materials for hydrogenation reaction.
The top of the nitric acid concentration tower 12 is mainly used for discharging acid-containing wastewater outside a boundary area through a pipeline 30 for environment-friendly treatment, concentrated nitric acid concentrated at the bottom of the nitric acid concentration tower enters an NO recovery tower 13 through a pipeline 31 to be used as an acid source, and recovered methanol from a pipeline 57 is in countercurrent contact with purge gas from a pipeline 32 to generate esterification regeneration reaction so as to recover NO in the purge gas; nitric acid waste liquid containing methanol at the tower bottom of the NO recovery tower 13 enters the methanol recovery tower 11 through a pipeline 33 for recycling, and light components containing MN generated at the tower top enter the MN recovery tower 15 after being pressurized by a compressor 14. In MN recovery tower 15, the recovered methanol from pipeline 29 is contacted in countercurrent, MN in the recovered methanol is eluted, and the eluted methanol enters esterification reaction tower 9 from tower bottom through pipeline 36, the gas phase light component at the tower top enters pressure swing adsorption tank 16 through pipeline 37, and CO is removed through pressure swing adsorption2The mixed gas containing CO is fed into the dehydration column 10 through a pipe 39, and the CO is removed2The gas can be discharged to the outside of the battery limits for treatment.
From the conduit 54Industrial hydrogen and circulating gas from a pipeline 53 are mixed, pressurized by a hydrogenation circulating compressor 25, enter a pipeline 55, then are mixed with dimethyl oxalate components from a pipeline 15 and recovered hydrogen from a pipeline 68 to serve as hydrogenation reaction raw materials, enter an outlet heat exchanger II 20 from a pipeline 40, perform heat exchange with hydrogenation reaction products led out from the bottom of a hydrogenation reactor 17, and then enter a catalyst bed layer from the top of the hydrogenation reactor 17 to perform catalytic hydrogenation reaction; meanwhile, refined water from the outside of the system enters the steam drum II 18 through a pipeline 48, a refrigerant in the steam drum II 18 enters a plate group fixing cavity from the bottom of the hydrogenation reactor 17 through a pipeline 49 to perform heat exchange with a catalyst bed layer, heat generated by the reaction is removed, the heated refrigerant is a vapor-liquid mixture, the vapor-liquid mixture is led out from the top of the hydrogenation reactor 17 and then enters the steam drum II 18 to perform gas-liquid separation, and generated low-pressure saturated steam enters an outside low-pressure steam recovery system through a pipeline 47 to be recycled. The hydrogenation reaction product enters the high-pressure gas-liquid separator 21 from the pipeline 44 after heat exchange for gas-liquid separation, most of the gas phase part passes through the pipeline 51 and then enters the pipeline 53 as circulating gas for circulation, and the rest part of the gas enters the low-pressure gas-liquid separator 26 through the pipeline 52 for gas-liquid separation; the liquid phase methanol in the low pressure gas-liquid separator 26 flows out through a pipeline 64, the gas phase part is merged with the noncondensable gas from the pipeline 58 through a pipeline 65 and then enters a methanol absorption tank 27 through a pipeline 66 to further remove the methanol, the gas after liquid removal enters a membrane separator 28 through a pipeline 67 and is subjected to recovery treatment of a membrane system to remove a small part of CO2CO and CH4When uncondensed steam is discharged from the pipeline 69, most of the recovered H2After being pressurized, the mixture enters a pipeline 68 for recycling.
The liquid-phase ethylene glycol crude product separated from the high-pressure gas-liquid separator 21 flows out from the pipeline 50, joins with liquid-phase methanol from the pipeline 64 and then enters the methanol separation tower 22; a certain amount of non-condensed steam is discharged from the top of the methanol separation tower 22 through a pipeline 58 for recovery, the light components of the liquid phase at the top of the tower enter a pipeline 57, and the liquid phase at the bottom of the tower enters a light component rectifying tower 23 through a pipeline 56 for separation; light components such as light components ethanol, methyl glycolate and the like at the top of the light component rectifying tower 23 enter a boundary region outside alcohol recovery device through a pipeline 60 for recovery, a tower bottom polyol mixture enters an ethylene glycol product tower 24 through a pipeline 59 for further purification, wherein the mixed light components mainly containing 1, 2-BDO and ethylene glycol are further recovered and treated through a pipeline 63, the ethylene glycol produced at the side line of the upper part of the tower body is taken out as a product through a pipeline 62, and the tower bottom is a mixture containing a small amount of ethylene glycol and ethylene glycol polycondensate and enters a boundary region outside for treatment.
At the initial start-up stage, the start-up heater 19 is used for heating the hydrogenation raw material, the heat source adopts low-pressure steam, the hydrogenation raw material from the pipeline 40 enters the pipeline 45, is preheated to the bed inlet temperature by the start-up heater 19, and then enters the catalyst bed from the top of the hydrogenation reactor 17 through the pipeline 46 and the pipeline 42 for hydrogenation reaction.
Examples of industrial production using the above process flow are as follows:
overhead light fraction from methanol washing column (composition: MN:5.22 v%, CO:22.12 v%, N)2：58.5v％，NO:11.14v％，CO20.63 v%, methanol 1.57 v%, and others 0.82 v%) and NO from outside the battery limits were mixed and fed into an esterification reaction tower 9 (inner diameter 50mm, height 2600mm, theoretical plate number 25, plate structure packed tower) from the 25 th plate, O2Feeding into esterification reaction tower 9 from 22 th, 23 th and 25 th trays respectively, and performing gas-liquid countercurrent contact with fresh methanol fed from tower top 1 st tray and recovered methanol mixed solution from methanol recovery tower 11, methanol and methyl nitrite mixture recovered from methanol rectification tower 5 fed from 5 th tray, methyl nitrite-containing alcohol solution from tower bottom of MN recovery tower 15 and tower bottom reflux liquid from tower bottom of 10 th tray in tower to generate esterification reaction (wherein O is2The molar ratio of NO to methanol is as follows: 0.1:0.6:50). The temperature of the top of the esterification reaction tower 9 is 50 ℃, the temperature of the bottom of the esterification reaction tower is 93 ℃, the temperature of the reaction zone is 70 +/-10 ℃, and the reaction pressure is 2 MPa. The bottom discharge of the esterification tower 9 (the composition is 71.8 wt% of methanol, 8.0 wt% of MN, and 20.2 wt% of other heavy components such as acid and water generated in the reaction) is extracted and then enters a methanol recovery tower 11 for recovery treatment. Gas phase at the top of esterification tower 9Component (composition: MN:10.05 v%, CO:26.42 v%, N)2：55.88v％，NO:5.2v％，CO20.60 v%, methanol 1.57 v%, and others 0.28 v%) are dehydrated in the dehydration tower 10. The dried gas with water content of 60ppm is obtained after dehydration in a dehydration tower 10 (the adsorbent is 4A molecular sieve, the operation temperature is 43 ℃, the pressure is 1.9MPa, and the two molecular sieve dryers A and B alternately operate and regenerate).
Acid-containing waste alcohol liquid in the tower kettle of the esterification reaction tower 9 enters a methanol recovery tower 11 (the inner diameter is 50mm, the height is 2100mm, the number of theoretical plates is 20, high-efficiency structured packing is filled in the methanol recovery tower 11, the temperature of the top of the tower is 120 ℃, the temperature of the bottom of the tower is 140 ℃, the pressure of the top of the tower is 0.7MPa, the reflux ratio of light components at the top of the tower is 1.2, and light components containing methanol (the components are 90 wt% of methanol, 8 wt% of MN and 8 wt% of22 wt%) of the methanol, and the rest of the methanol is taken as the washing liquid in the MN recovery tower 15; the acid-containing wastewater at the tower bottom of the methanol recovery tower 11 enters a nitric acid concentration tower 12 for nitric acid concentration.
The carbonylation reactor 1 (plate type fixed bed reactor, the inner diameter is 320mm, the height is 2000mm), the centre has fixed cavity of the slab group, there are 3 groups of slabs in the fixed cavity of the slab group, each group has 3 slabs; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the carbonylation reactor 1, and a carbonylation high-pressure reaction catalyst (a catalyst sold by Shanghai Peng Zheng engineering technology Co., Ltd., and the commercial brand of the catalyst is DMO-0701T) is filled in the catalyst bed layer. The dry gas from the dehydration tower 10 is mixed with dehydrogenation-treated industrial grade CO (99 v%) as a carbonylation reaction raw material and nitrogen as an inert gas source, then the mixture is subjected to heat exchange with a carbonylation reaction product through an outlet heat exchanger I3, the mixture is preheated to 95 ℃, the mixture firstly enters from the top of a carbonylation reactor 1 and then enters a catalyst bed layer in a radial flow mode to carry out carbonylation reaction (the hot spot temperature of the catalyst bed layer is 130 ℃, the reaction pressure is 1.8MPa, and the gas hourly space velocity is 10000h-1) (ii) a The carbonylation product enters an outlet heat exchanger 3 for heat exchange and then enters a first gas-liquid separator 4 for gas-liquid separation.
The coolant of the fixed cavity of the plate set of the carbonylation reactor 1 is water medium. Refined water from the outside of the system enters a steam pocket I2 to supplement water, the water in the steam pocket I enters a plate group fixing cavity in the carbonylation reactor 1 to exchange heat with a catalyst bed layer, the heat generated by the reaction is removed, the heated water is a vapor-liquid mixture and enters the steam pocket to be subjected to gas-liquid separation, and the generated low-pressure saturated steam is sent to a low-pressure steam pipe network outside a boundary area to be recycled.
The liquid phase (methanol: 1.16 wt%, DMC: 0.45 wt%, DMO: 97.6 wt%, and other 0.79 wt%) from the first gas-liquid separator 4 is fed into a methanol rectifying tower 5 as an extractant for separation; leading out the mixed gas phase component containing DMC into a methanol washing tower 7 (the inner diameter is 50mm, the height is 3200mm, the number of theoretical plates is 30, high-efficiency structured packing is arranged in the methanol washing tower 7, the temperature of the top of the tower is 28.1 ℃, the temperature of a tower kettle is 39.8 ℃, and the pressure of the top of the tower is 1.5MPa), and the mixed gas containing DMC and DMO are eluted by countercurrent contact with the recovered methanol (the content is 99.9 wt%) in a methanol separation tower 22, most of the gas phase light component at the top of the methanol washing tower 7 enters an esterification reaction tower 9 through a carbonylation circulating compressor 8, and the nitrogen oxide generated by carbonylation reaction is recycled; a small part of non-condensable gas (gas accounts for 0.5 v%) is used as purge gas and enters an NO recovery tower 13 for recovery treatment; the liquid phase in the bottom of the methanol washing tower 7 enters a methanol rectifying tower 5 for separation.
Methanol rectifying tower 5 (inner diameter: 50mm, height: 2600mm, extractive rectifying tower, theoretical plate number: 25, packed with highly efficient structured packing, overhead temperature of 73.12 deg.C, bottom temperature of 185.0 deg.C, top pressure of 0.1MPa) light components (methanol: 88.2 wt%, MN: 11.8 wt%) enter esterification tower 9 as one of alcohol sources, and bottom contains
Heavy components of DMC and DMO enter a DMO rectifying tower 6 for separation.
DMO rectifying tower 6 (inner diameter: 50mm, height 3000mm, theoretical plate number 28, packed with high efficiency structured packing, tower top temperature 103 deg.C, tower bottom temperature 180 deg.C, normal pressure operation, reflux ratio 50), and collecting DMC as product at tower top (DMC product purity is 99.41 wt%); the heavy components in the tower bottom (DMO purity is 99.9 wt%) are all used as raw materials of a hydrogenation section.
In the nitric acid concentration tower 12 (the inner diameter is 32mm, the height is 850mm, the number of theoretical plates is 8, high-efficiency structured packing is filled in the nitric acid concentration tower 12, the temperature of the top of the tower is 64 ℃, the temperature of a tower kettle is 87 ℃, the pressure of the top of the tower is 0.15MPa, and the reflux ratio is 0.05), the top of the tower is mainly used for discharging acid-containing wastewater outside a boundary area for environment-friendly treatment, and concentrated nitric acid with the concentration of 68 wt% is generated by concentration of the tower kettle and is used as an acid source of.
In an NO recovery tower 13 (inner diameter: 32mm, height: 2100mm, theoretical plate number: 20, packed with high-efficiency structured packing, overhead temperature of 50 ℃, bottom temperature of 100 ℃ and overhead pressure of 1.4MPa), the purge gas from the methanol washing tower 7 is fed from the 20 th tray, the recovered methanol (99.9 wt%) from the methanol separation tower 22 fed from the 1 st tray and the concentrated nitric acid from the nitric acid concentration tower 12 fed from the 8 th tray are contacted in a countercurrent manner to carry out esterification regeneration reaction. NO in the purge gas and HNO in concentrated nitric acid3The molar ratio of methanol was 1:2.5: 20. Light components (composition: CO:21.1 v%, CO) at the top of NO recovery column 132:0.6v％、MN:20.8v％、N2: 55.7 v%, methanol: 1.8 v%) is pressurized by a compressor 14 and enters an MN recovery tower 15; heavy components (the composition: methanol is 71.8 wt%, and other heavy components such as acid and water generated by the reaction are 28.2 wt%) in the tower bottom of the NO recovery tower 13 enter a tower plate 3 of the methanol recovery tower 11 for recovery.
The feed in MN recovery column 15 (inner diameter: 32mm, height: 3200mm, theoretical plate number: 30, packed with high-efficiency structured packing, overhead temperature: 30.8 ℃, column bottom temperature: 41.3 ℃ and overhead pressure: 2MPa) was brought into countercurrent contact with the recovered methanol from methanol recovery column 11 fed from tray 1 to absorb a large amount of MN in the feed gas and the rest gas (composition: CO:27.3 v%, CO content: balance)2:0.8v％，N2: 71.9 v%,) from the top of the column into a pressure swing adsorption tank 16, and the contents of the bottom of the column (composition: methanol: 79.3 mol%, MN: 20.7 mol%) is recycled by entering the 5 th tray of the esterification reaction tower 9. The gas phase at the top of the MN recovery tower 15 is subjected to pressure swing adsorption by a pressure swing adsorption tank 16, and the purified gas (N)2: 72 v%, CO: 28 v%) into the dehydration tower 10 for treatmentCarbonylation reactor 1, and 0.95 v% gas (composition: CO)299.8 v%) was discharged outside the battery limits for treatment.
A hydrogenation reactor 17 (a plate type fixed bed hydrogenation reactor, the inner diameter is 325mm, the height is 900mm), a plate group fixing cavity is arranged at the center, three groups of plates are arranged in the plate group fixing cavity, and each group comprises 3 plates; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the hydrogenation reactor, and is filled with a hydrogenation reaction catalyst: the catalyst is sold in Shanghai Peng engineering technology Co., Ltd, and the trade name of the catalyst is MEG-801T).
Technical grade H2(the purity is 99.9 v%) and the recycle gas from the high-pressure gas-liquid separator 21 (the composition is that hydrogen is 96 v%, methane is 0.05 v%, nitrogen is 0.02 v%, carbon monoxide is 0.02 v%, methanol is 3 v%, and other 0.91 v%) are compressed by a hydrogenation recycle compressor 25, then are merged with dimethyl oxalate (99.9 wt%) from the tower bottom of a DMO rectifying tower 6, enter an outlet heat exchanger II 20 of a hydrogenation plate reactor 17, are preheated to 175 ℃, firstly enter from the top of the hydrogenation reactor 17, and then enter a catalyst bed layer in a radial flow mode to carry out hydrogenation reaction (the hot spot temperature of the catalyst bed layer is 190 ℃, the reaction pressure is 3.0MPa, and the liquid hourly space velocity is 2.8 Kg/Kg.h); and the hydrogenated product is discharged from the bottom, enters an outlet heat exchanger II 20 for heat exchange, and then enters a high-pressure gas-liquid separator 21 for gas-liquid separation.
At the initial stage of start-up, the material passing through the outlet heat exchanger II 20 enters the start-up heater 19 for preheating, and the preheated gas serving as the feed gas enters the catalyst bed layer for hydrogenation after reaching the inlet temperature of the catalyst bed layer.
The refrigerant of the fixed cavity of the sheet group of the hydrogenation reactor 17 is an aqueous medium. Refined water from the outside of the system enters a steam drum II 18 to supplement water, water in the steam drum II 18 enters a plate group fixing cavity in a hydrogenation reactor 17 to exchange heat with a catalyst bed layer, heat generated by reaction is removed, the heated water is a steam-liquid mixture and enters the steam drum II to be subjected to gas-liquid separation, and generated low-pressure saturated steam is sent to a low-pressure steam pipe network outside a boundary area to be recycled.
After the hydrogenation product is separated by the high-pressure gas-liquid separator 21, most of the gas phase enters a hydrogenation circulating compressor 25 as the circulating gas, the rest of the non-condensable gas (gas content: 1.2 v%) enters a low-pressure gas-liquid separator 26, and the liquid phase (methanol: 50.1 wt%, ethylene glycol: 48.55 wt%, methyl glycolate: 0.06 wt%, ethanol: 0.39 wt%, BDO:0.12 wt%, and the rest: 0.78 wt%) led out from the high-pressure gas-liquid separator 21 enters a methanol separation tower 22 for separation. The liquid phase separated by the low-pressure gas-liquid separator 26 enters a methanol separation tower 22 for separation, and after the gas phase is further methanol-removed by a methanol absorption tank 27 (with the inner diameter of 160mm and the height of 900mm), the gas phase (comprising 97 v% of hydrogen, 0.15 v% of methane, 0.06 v% of nitrogen, 0.27 v% of carbon monoxide and the rest 2.52 v%) enters a membrane separator 28 for recycling. The hydrogen (with the purity of 99.9 v%) separated by the membrane separator is preheated by the outlet heat exchanger II and then enters the hydrogenation plate type reactor 17, and only a small part of non-condensable gas such as rich methane is discharged as purge gas to the outside for recycling.
In a methanol separation tower 22 (the inner diameter is 50mm, the height is 2600mm, the number of theoretical plates is 25, efficient structured packing is filled in the tower, the temperature of the tower top is 50.82 ℃, the temperature of a tower kettle is 171 ℃, and the absolute pressure of the tower top is 90kPa), materials are fed at a 12 th tower plate, non-condensable gas at the tower top enters a methanol absorption tank 27 for treatment and then enters a membrane separator 28, the reflux ratio at the tower top is 1.6, and discharged materials (99.9 wt% of methanol and 0.1 wt% of other low boiling point components) at the tower top are extracted and then respectively enter a methanol washing tower 7 and an NO recovery tower 13; heavy components (composed of 96 wt% of ethylene glycol, 0.12 wt% of methyl glycolate, 2.68 wt% of 1.2-BDO, 0.8 wt% of ethanol and 0.4 wt% of other components) in the bottom of the methanol separation tower 22 enter a light component rectifying tower 23.
A light component rectifying tower 23 (the inner diameter is 50mm, the height is 4000mm, the number of theoretical plates is 40, efficient structured packing is filled in the rectifying tower, the temperature of the tower top is 83.8 ℃, the temperature of a tower kettle is 146.9 ℃, the absolute pressure of the tower top is 16kPa, the reflux ratio of the tower top is 50), and an ethanol crude product (98 wt% ethanol and 2 wt% methyl glycolate) is led out of the tower top and sent to a boundary region for collection treatment; heavy components in the bottom of the column (97.9 wt% ethylene glycol, 2.1 wt% 1.2-BDO) ethylene glycol product column 24.
In an ethylene glycol product tower 24 (the inner diameter is 50mm, the height is 6500mm, the theoretical plate number of the tower is 60, high-efficiency structured packing is filled in the tower, the temperature of the top of the tower is 130 ℃, the temperature of a tower bottom is 170.1 ℃, and the absolute pressure of the top of the tower is 5kPa), the reflux ratio of the top of the tower is 98, the components of 1, 2-BDO (19.79 wt percent, 80wt percent of ethylene glycol and 0.21wt percent of the ethylene glycol) are extracted from the top of the tower to the outside of a battery limits and recovered as a byproduct, the condensation polymer of the ethylene glycol and the ethylene glycol is treated outside the battery limits in the tower bottom, and the final product of the ethylene glycol (the content is 99.99wt percent) is extracted from.
The foregoing embodiments are merely illustrative of the principles and utilities of the present invention and are not intended to limit the invention. Any person skilled in the art can modify or change the above-mentioned embodiments without departing from the spirit and scope of the present invention. Accordingly, it is intended that all equivalent modifications or changes which can be made by those skilled in the art without departing from the spirit and technical spirit of the present invention be covered by the claims of the present invention.
Claims (42)
1. A process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation specifically comprises the following steps:
(1) introducing industrial grade NO and O into an esterification reaction tower (9)2Carrying out esterification reaction with methanol; introducing the methyl nitrite mixed gas at the tower top of the esterification reaction tower (9) into a carbonylation reactor (1) for carbonylation reaction; part of the acidic alcohol solution at the bottom of the esterification tower (9) reflows to the esterification tower (9), and part of the acidic alcohol solution is introduced into a methanol recovery tower (11); the methanol recovered from the top of the methanol recovery tower (11) is partially recycled to the esterification reaction tower (9) for recycling, and the methanol is recycledThe rest enters an MN recovery tower (15) to be used as washing liquid; waste acid at the tower bottom of the methanol recovery tower (11) enters a nitric acid concentration tower (12) for concentration treatment;
(2) methyl nitrite from esterification reaction tower and industrial grade CO and N2Feeding the material into a carbonylation reactor (1) to carry out carbonylation reaction in the presence of a carbonylation reaction catalyst; the temperature of the carbonylation reaction is 30-200 ℃, the reaction pressure is 1-10 MPa, and the gas hourly space velocity is 3000-30000 h-1；
(3) The carbonylation product enters a first gas-liquid separator (4) for gas-liquid separation, the gas phase enters a methanol washing tower (7), and the liquid phase enters a methanol rectifying tower (5); the gas phase component at the top of the methanol washing tower (7) is partially circulated to the esterification reaction tower (9), and part of the gas phase component is used as purge gas to enter an NO recovery tower (13) for recovery treatment; the liquid phase component of the tower bottom of the methanol washing tower (7) enters a methanol rectifying tower (5) for rectification and separation; the mixture of methanol and methyl nitrite recovered from the top of the methanol rectifying tower (5) is recycled to the esterification reaction tower (9) for reuse, and heavy components in the tower bottom enter the DMO rectifying tower (6); obtaining a DMC product at the top of the DMO rectifying tower (6), and enabling dimethyl oxalate components at the bottom of the tower to enter a hydrogenation reactor (17) for hydrogenation reaction;
(4) waste acid from the methanol recovery tower (11) is concentrated by a nitric acid concentration tower (12) until the concentration of nitric acid is 10-68 wt%, and then is circulated to an NO recovery tower (13); in the NO recovery tower (13), concentrated nitric acid, methanol and purge gas from a methanol washing tower (7) are subjected to esterification regeneration reaction; the gas phase light component at the top of the NO recovery tower (13) enters an MN recovery tower (15), and the nitric acid waste liquid containing methanol generated at the bottom of the tower is circulated to a methanol recovery tower (11) for further recovery treatment; in the MN recovery tower (15), gas phase feed enters a pressure swing adsorption tank (16) after being washed by recovered methanol, and an alcoholic solution containing methyl nitrite in the tower kettle of the MN recovery tower (15) circularly enters an esterification reaction tower (9); CO separated from the pressure swing adsorption tank (16)2Discharging to outside for treatment, recovering N2And the purified CO gas enters the carbonylation reactor (1) for cyclic utilization;
(5) dimethyl oxalate components from the tower bottom of the DMO rectifying tower (6) and industrial hydrogen pressurized by a hydrogenation circulating compressor (25) are mixed and then enter a hydrogenation reactor (17), and in the presence of a hydrogenation catalyst, hydrogenation reaction is carried out to generate methanol, ethylene glycol and the like; the hydrogenation reaction temperature is 160-320 ℃, the reaction pressure is 1-10 MP, and the liquid hourly space velocity is 1-3 Kg/Kg.h;
(6) and the hydrogenated product enters a second gas-liquid separator for gas-liquid separation, a gas phase part is pressurized by the hydrogenation circulating compressor (25) and then circulates to the hydrogenation reactor (17), a part of the gas phase enters a membrane separator (28), the gas phase is recycled by returning to the hydrogenation reactor (17), and a liquid phase enters an ethylene glycol product tower (24) for separation to obtain an ethylene glycol product.
2. The process for producing dimethyl oxalate and hydrogenating to ethylene glycol through high pressure carbonylation of industrial synthesis gas according to claim 1, further comprising any one or more of the following features:
the carbonylation reactor (1) is externally connected with a dehydration tower (10); the gas phase recovered by the pressure swing adsorption tank (16) and the methyl nitrite mixed gas from the top of the esterification reaction tower (9) are dehydrated by the dehydrating tower (10) and then enter the carbonylation reactor (1) for carbonylation reaction;
(II) an outlet heat exchanger I (3) is connected outside the carbonylation reactor (1); industrial grade CO, N2And dry gas from the dehydrating tower (10) is used as a carbonylation reaction raw material, exchanges heat with a carbonylation reaction product from the carbonylation reactor (1) through the outlet heat exchanger I (3), and then enters the carbonylation reactor (1) for carbonylation reaction;
thirdly, pressurizing part of the gas phase component from the top of the methanol washing tower (7) by a carbonylation circulating compressor (8) and then entering an esterification reaction tower (9);
(IV) an outlet heat exchanger II (20) is connected outside the hydrogenation reactor (17); dimethyl oxalate components from a DMO rectifying tower (6), industrial hydrogen and recycle gas from a pressurized recycle compressor and recycle gas from a membrane separator (28) are used as hydrogenation reaction raw materials, and the raw materials and the hydrogenation products from the hydrogenation reactor (17) enter the hydrogenation reactor (17) for hydrogenation reaction after heat exchange through the outlet heat exchanger II (20);
fifthly, the gas phase light component at the top of the NO recovery tower (13) enters an MN recovery tower (15) after being compressed and pressurized by a compressor (14);
(sixth) the liquid phase separated by the second gas-liquid separator is firstly fed into a methanol separation column (22); the non-condensable gas recovered from the top of the methanol separation tower (22) enters the membrane separator (28), part of liquid phase light components such as methanol recovered from the top of the methanol separation tower (22) enters the upper part of the methanol washing tower (7) to be used as washing liquid, and part of the liquid phase light components enters the NO recovery tower (13); the heavy component of the liquid phase at the bottom of the methanol separation tower (22) enters a light component rectifying tower (23) for further separation and purification; the light components at the top of the light component rectifying tower (23) enter an out-of-range alcohol recovery device for recovery treatment; heavy components in the bottom of the light component rectifying tower (23) enter the ethylene glycol product tower (24); light components at the top of the ethylene glycol product tower (24) enter an outside-of-the-road 1, 2-BDO recovery processing device for further recovery processing, heavy components at the bottom of the ethylene glycol product tower (24) enter the outside-of-the-road recovery processing device for subsequent processing, and an ethylene glycol product is led out from the upper side line of the ethylene glycol product tower (24).
3. The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation according to claim 2, wherein the dehydration tower comprises two molecular sieve dryers A and B which alternately operate and regenerate; the molecular sieve dryer A and the molecular sieve dryer B are filled with adsorbents; the adsorbent is selected from a 3A molecular sieve, a 4A molecular sieve, a 5A molecular sieve, a 9A molecular sieve and calcium oxide; the operation temperature of the molecular sieve dryer A and the molecular sieve dryer B is 40-260 ℃, and the pressure is 1-10 MPa; the drying gas is obtained after being treated by a dehydrating tower (10), and the water content in the drying gas is 0.1-100 ppm.
4. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 2, wherein the second gas-liquid separator comprises a high-pressure gas-liquid separator (21) and a low-pressure gas-liquid separator (26); the gas phase separated by the high-pressure gas-liquid separator (21) enters the hydrogenation circulating compressor (25), and the gas phase enters the low-pressure gas-liquid separator (26); the liquid phase separated by the high-pressure gas-liquid separator (21) enters the methanol separation tower (22); the gas phase separated by the low-pressure gas-liquid separator (26) enters the membrane separator (28), and the liquid phase separated by the low-pressure gas-liquid separator (26) enters the methanol separation tower (22).
5. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation as claimed in claim 4, wherein 0.1-10 v% of gas phase separated by the high-pressure gas-liquid separator enters the low-pressure gas-liquid separator.
6. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 4, wherein the gas phase separated by the low-pressure gas-liquid separator (26) and the noncondensable gas from the top of the methanol separation tower (22) enter the membrane separator (28) after methanol is absorbed by the methanol absorption tank (27).
7. The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and hydrogenating to produce ethylene glycol as claimed in claim 1, further comprising any one or both of the following features:
the carbonylation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor;
and (II) the hydrogenation reactor is a plate reactor, a tubular reactor or a tubular-plate composite reactor.
8. The process for producing dimethyl oxalate and hydrogenating to ethylene glycol through high pressure carbonylation of industrial synthesis gas according to claim 7, further comprising any one or both of the following features:
the carbonylation reactor (1) is a plate type fixed bed carbonylation reactor; a plate group fixing cavity is arranged in the center of the plate type fixed bed carbonylation reactor, and a plate group is arranged in the plate group fixing cavity; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed carbonylation reactor; the catalyst bed layer is filled with a carbonylation reaction catalyst; after the carbonylation reaction raw material reaches the inlet temperature of the catalyst bed layer, the carbonylation reaction raw material enters the catalyst bed layer from the top of the plate type fixed bed carbonylation reactor to carry out carbonylation reaction; the method comprises the following steps that a refrigerant introduced from the outside enters a plate group fixing cavity from the bottom of a plate type fixed bed carbonylation reactor and is led out from the top of the plate type fixed bed carbonylation reactor, and the heat exchange is carried out in the countercurrent process to take away the reaction heat of the carbonylation reaction; the carbonylation product from the bottom of the catalyst bed is led out from the bottom of the plate type fixed bed carbonylation reactor;
(II) the hydrogenation reactor (17) is a plate-type fixed bed hydrogenation reactor; a plate group fixing cavity is arranged in the center of the plate type fixed bed hydrogenation reactor, and a plate group is arranged in the plate group fixing cavity; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed hydrogenation reactor; the catalyst bed layer is filled with a hydrogenation reaction catalyst; after the hydrogenation reaction raw material reaches the inlet temperature of the catalyst bed, the hydrogenation reaction raw material enters the catalyst bed from the top of the plate-type fixed bed hydrogenation reactor to carry out hydrogenation reaction; a refrigerant introduced from the outside enters the plate group fixing cavity from the bottom of the plate type fixed bed hydrogenation reactor and is led out from the top of the plate type fixed bed hydrogenation reactor, and the heat exchange is carried out in the countercurrent process to take away the reaction heat of the hydrogenation reaction; the hydrogenation product from the bottom of the catalyst bed is drawn from the bottom of the plate fixed bed hydrogenation reactor.
9. The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation as claimed in claim 8, wherein a steam drum I (2) is connected outside the plate-type fixed bed carbonylation reactor; the method comprises the following steps that (1) a refrigerant introduced from the outside enters a steam pocket I (2), the refrigerant in the steam pocket I (2) enters a plate group fixing cavity of a plate type fixed bed carbonylation reactor to exchange heat with a catalyst bed layer, and reaction heat is removed; the heated refrigerant is a vapor-liquid mixture, enters a steam drum I (2) for gas-liquid separation, and the generated low-pressure saturated steam enters an outdoor low-pressure steam recovery system for recycling; a steam drum II (18) is connected outside the plate-type fixed bed hydrogenation reactor; the method comprises the following steps that (1) a refrigerant introduced from the outside enters a steam pocket II (18), the refrigerant in the steam pocket II (18) enters a plate group fixing cavity of a plate type fixed bed hydrogenation reactor to exchange heat with a catalyst bed layer, and reaction heat is removed; the heated refrigerant is a vapor-liquid mixture, enters a steam drum II (18) for gas-liquid separation, and the generated low-pressure saturated steam enters an outdoor low-pressure steam recovery system for recycling.
10. The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation according to claim 8, wherein a start-up heater (19) is connected outside the plate-type fixed bed hydrogenation reactor; in the initial startup stage, the hydrogenation reaction raw material comes out of the outlet heat exchanger II (20) and then enters the startup heater (19) for preheating, and enters the catalyst bed layer for hydrogenation reaction after the preheating reaches the inlet temperature of the catalyst bed layer; in the initial startup stage, the startup heater (19) provides a unique heat source for the hydrogenation reaction in the plate-type fixed bed hydrogenation reactor; the heat source of the start-up heater (19) is low-pressure steam.
11. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the esterification reaction tower (9) is 20-50; in the feed of the esterification reaction tower (9), the O2Feeding from 16 th to 50 th tower plates respectively in 2-8 paths; the NO and the overhead gas phase light components from the methanol washing tower (7) are fed from 18 th to 50 th tower plates; feeding the fresh methanol, the recovered methanol from the top of the methanol recovery tower (11), the methanol and methyl nitrite mixture recovered from the top of the methanol rectifying tower (5) and the methyl nitrite-containing alcohol solution from the tower kettle of the MN recovery tower (15) from tower plates 1-5; feeding reflux materials in a tower kettle of the esterification reaction tower (9) from a 10 th to a 25 th tower plate;
in the esterification reaction tower (9), O2The molar ratio of NO to methanol is 0.01-0.8: 0.1-3.2: 0.8 to 50; the temperature of the top of the esterification reaction tower (9) is 30-80 ℃, the temperature of the tower kettle is 50-200 ℃, the temperature of the reaction zone is 50-160 ℃, and the reaction pressure is 0.5-10 MPa。
12. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the methanol recovery tower (11) is 5-50, the temperature of the top of the tower is 40-150 ℃, the temperature of the bottom of the tower is 60-230 ℃, and the pressure of the top of the tower is 0.1-2.0 MPa; the reflux ratio of the light components at the top of the methanol recovery tower (11) is 0.1-3.0; in the methanol recovery tower, the part which circularly enters the esterification reaction tower accounts for 10-90 wt% of the recovered methanol at the top of the methanol recovery tower.
13. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the methanol washing tower (7) is 10-50, the temperature at the top of the tower is 15-70 ℃, the temperature at the bottom of the tower is 10-100 ℃, and the pressure at the top of the tower is 0.9-10 MPa; in the gas phase component at the top of the methanol washing tower, the proportion of the purge gas is 0.05-5 v%.
14. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the methanol rectifying tower (5) is an extractive rectifying tower, the number of theoretical plates is 10-60, the temperature at the top of the tower is 50-150 ℃, the temperature at the bottom of the tower is 130-250 ℃, and the pressure at the top of the tower is 0.01-0.5 MPa.
15. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the nitric acid concentration tower (12) is 1-30, the temperature of the top of the tower is 30-110 ℃, the temperature of the bottom of the tower is 60-160 ℃, and the pressure of the top of the tower is 0.01-0.3 MPa; the reflux ratio of the light components at the top of the nitric acid concentration tower (12) is 0.01-3.
16. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the NO recovery tower (13) is 5-30, the temperature at the top of the tower is 30-120 ℃, the temperature at the bottom of the tower is 50-200 ℃, and the pressure at the top of the tower is 1-10 MPa; feeding the purge gas from a tower plate 5-30 of an NO recovery tower (13); feeding the concentrated nitric acid from a 1 st to a 10 th tower plate of an NO recovery tower (13); feeding recovered methanol from the top of the methanol separation tower (22) from tower plates 1-10; in the NO recovery tower (13), the molar ratio of the nitric acid to the methanol to the NO in the purge gas is 1.1-10: 2-100: 1-5.
17. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the MN recovery tower (15) is 10-60, the temperature of the top of the tower is 0-50 ℃, the temperature of the bottom of the tower is 0-80 ℃, and the reaction pressure is 1-10 MPa.
18. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 1, wherein the number of theoretical plates of the DMO rectifying tower (6) is 10-50, the temperature of the top of the tower is 80-120 ℃, the temperature of the bottom of the tower is 120-200 ℃, and the process is operated under normal pressure or reduced pressure; the reflux ratio of light components at the top of the DMO rectifying tower (6) is 0.1-100.
19. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 2, wherein the number of theoretical plates of the methanol separation tower (22) is 10-40, the temperature of the top of the tower is 40-70 ℃, the temperature of the bottom of the tower is 80-180 ℃, and the operation is carried out under normal pressure or reduced pressure; the reflux ratio of light components at the top of the methanol separation tower is 0.1-3.
20. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 2, wherein the number of theoretical plates of the light component rectifying tower (23) is 10-60, the temperature of the top of the tower is 58-90 ℃, the temperature of the bottom of the tower is 70-160 ℃, and the absolute pressure of the top of the tower is 5-50 KPa; the reflux ratio of the light components at the top of the light component rectifying tower (23) is 1-50.
21. The process for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 2, wherein the number of theoretical plates of the ethylene glycol product tower (24) is 30-100, the temperature of the top of the tower is 100-150 ℃, the temperature of the bottom of the tower is 130-230 ℃, and the absolute pressure of the top of the tower is 5-50 KPa; the reflux ratio of light components at the top of the ethylene glycol product tower is 50-200 or total reflux.
22. The process of claim 1, wherein the purified gas recovered from the pressure swing adsorption tank comprises the following components: n is a radical of260-80 v% of CO, and 20-40 v% of CO; separated CO2The gas accounts for 0.1-5 v% of the total amount of the intake air, wherein CO2The concentration of (A) is 99.8-99.9 v%.
23. The process for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation according to claim 2, wherein the membrane separator (28) is composed of a plurality of hollow fiber membrane modules which are connected in parallel or in series; the concentration of hydrogen in the purified gas obtained by separation and purification of the membrane separator (28) is 88-99.99 v%, and the recovery rate of hydrogen is 90-98.5%.
24. A device system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation is characterized by comprising a carbonylation reaction system, an esterification reaction system, a purge gas and waste acid coupling recovery system and a hydrogenation reaction system;
the carbonylation reaction system comprises a carbonylation reactor (1), a first gas-liquid separator (4), a methanol washing tower (7), a methanol rectifying tower (5) and a DMO rectifying tower (6); the carbonylation reactor (1) is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the first gas-liquid separator (4) is provided with a feed inlet, a gas outlet and a liquid outlet; the methanol washing tower (7) is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the methanol rectifying tower (5) is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the DMO rectifying tower (6) is provided with a lower feed inlet, a top outlet and a bottom outlet;
the esterification reaction system comprises an esterification reaction tower (9) and a methanol recovery tower (11); the esterification reaction tower (9) is provided with a top feed inlet, an upper feed inlet, a plurality of lower feed inlets, a middle reflux inlet, a top outlet and a bottom outlet; the methanol recovery tower (11) is provided with a middle-lower part feed inlet, a top outlet and a bottom outlet;
the purge gas and waste acid coupling recovery system comprises a nitric acid concentration tower (12), an NO recovery tower (13), an MN recovery tower (15) and a pressure swing adsorption tank (16); the nitric acid concentration tower (12) is provided with a middle feeding hole, a top outlet and a bottom outlet; the NO recovery tower (13) is provided with a top feed inlet, a middle feed inlet, a bottom feed inlet, a top outlet and a bottom outlet; the MN recovery tower (15) is provided with an upper feed inlet, a lower feed inlet, a top outlet and a bottom outlet; the pressure swing adsorption tank (16) is provided with a feed inlet, a recovered gas outlet and an exhaust gas outlet;
the hydrogenation reaction system comprises a hydrogenation circulating compressor (14), a hydrogenation reactor (17), a second gas-liquid separator, a membrane separator (28), a methanol separation tower (22), a light component rectifying tower (23) and an ethylene glycol product tower (24); the hydrocycle compressor (14) comprising an inlet and an outlet; the hydrogenation reactor (17) is provided with a top feeding hole, a bottom discharging hole, a bottom refrigerant inlet and a top refrigerant outlet; the second gas-liquid separator is provided with a feed inlet, a gas outlet and a liquid outlet; the membrane separator (28) is provided with a feed inlet, a recovered gas outlet and a discharged gas outlet; the methanol separation tower (22) is provided with a middle feed inlet, a top noncondensable gas outlet, a top liquid phase light component outlet and a bottom liquid phase heavy component outlet; the light component rectifying tower (23) is provided with a lower feed inlet, a top outlet and a bottom outlet; the ethylene glycol product tower (24) is provided with a lower feed inlet, a top outlet, an upper outlet and a bottom outlet;
the top feed inlet of the carbonylation reactor (1) is communicated with a CO raw material pipeline and N2The raw material pipelines are connected through pipelines; a discharge hole at the bottom of the carbonylation reactor (1) is connected with a feed hole of the first gas-liquid separator (4) through a pipeline; the gas outlet of the first gas-liquid separator (4) is connected with the lower feed inlet of the methanol washing tower (7) through a pipeline; a liquid outlet of the first gas-liquid separator (4) is connected with an upper feed inlet of the methanol rectifying tower (5) through a pipeline; a branch outlet A and a branch outlet B are arranged at the top outlet of the methanol washing tower (7), the branch outlet A is connected with a lower feed inlet of the esterification reaction tower (9) through a pipeline, and the branch outlet B is connected with a bottom feed inlet of the NO recovery tower (13) through a pipeline; the bottom outlet of the methanol washing tower (7) is connected with the lower feed inlet of the methanol rectifying tower (5) through a pipeline; the top outlet of the methanol rectifying tower (5) is connected with the upper feed inlet of the esterification reaction tower (9) through a pipeline; the bottom outlet of the methanol rectifying tower (5) is connected with the lower feed inlet of the DMO rectifying tower (6) through a pipeline; the bottom outlet of the DMO rectifying tower (6) is connected with the top feed inlet of the hydrogenation reactor (17) through a pipeline, and the top outlet of the DMO rectifying tower (6) is a DMC product outlet;
other lower feed inlets of the esterification reaction tower (9), an NO raw material pipeline and a multi-path O2The raw material pipelines are respectively connected through pipelines; a feed inlet at the top of the esterification reaction tower (9) is connected with a methanol raw material pipeline through a pipeline; a branch outlet C and a branch outlet D are arranged at the bottom outlet of the esterification tower (9), the branch outlet C is connected with the middle reflux inlet of the esterification tower (9) through a pipeline, and the branch outlet D is connected with the lower feed inlet of the methanol recovery tower (11) through a pipeline; the top outlet of the esterification reaction tower (9) is connected with the top feed inlet of the carbonylation reactor (1) through a pipeline; a branch outlet E and a branch outlet F are arranged at the top outlet of the methanol recovery tower (11), the branch outlet E is connected with the upper feed inlet of the esterification reaction tower (9) through a pipeline, and the branch outlet F is connected with the upper feed inlet of the MN recovery tower (15) through a pipeline; the bottom of the methanol recovery tower (11)The outlet of the part is connected with a middle feed inlet of the nitric acid concentration tower (12) through a pipeline;
an outlet at the top of the nitric acid concentration tower (12) is a waste liquid outlet; the bottom outlet of the nitric acid concentration tower (12) is connected with the middle feed inlet of the NO recovery tower (13) through a pipeline; the top outlet of the NO recovery tower (13) is connected with the lower feed inlet of the MN recovery tower (15) through a pipeline; the bottom outlet of the NO recovery tower (13) is connected with the middle-lower feed inlet of the methanol recovery tower (11) through a pipeline; the top outlet of the MN recovery tower (15) is connected with the feed inlet of the pressure swing adsorption tank (16) through a pipeline; the bottom outlet of the MN recovery tower (15) is connected with the upper feed inlet of the esterification reaction tower (9) through a pipeline; a recycled gas outlet of the pressure swing adsorption tank (16) is connected with a top feed inlet of the carbonylation reactor (1) through a pipeline; the exhaust gas outlet of the pressure swing adsorption tank (16) is connected with an out-of-range recovery device through a pipeline;
the inlet of the hydrogenation circulating compressor (14) is connected with an industrial hydrogen raw material pipeline through a pipeline, and the outlet of the hydrogenation circulating compressor (14) is connected with the top feed inlet of the hydrogenation reactor (17) through a pipeline; a discharge hole at the bottom of the hydrogenation reactor (17) is connected with a feed hole of the second gas-liquid separator through a pipeline; a gas outlet of the second gas-liquid separator is provided with a branch outlet G and a branch outlet H, the branch outlet G is connected with an inlet of the hydrogenation circulating compressor (14) through a pipeline, and the branch outlet H is connected with a feed inlet of the membrane separator (28) through a pipeline; the liquid outlet of the second gas-liquid separator is connected with the lower feed inlet of the methanol separation tower (22) through a pipeline; the top noncondensable gas outlet of the methanol separation tower (22) is connected with the feed inlet of the membrane separator (28) through a pipeline; a top liquid phase light component outlet of the methanol separation tower (22) is provided with a branch outlet I and a branch outlet J, the branch outlet I is connected with an upper feed inlet of the methanol washing tower (7) through a pipeline, and the branch outlet J is connected with a top feed inlet of the NO recovery tower (13) through a pipeline; a bottom liquid phase heavy component outlet of the methanol separation tower (22) is connected with a lower feed inlet of the light component rectifying tower (23) through a pipeline; the top light component outlet of the light component rectifying tower (23) is connected with an out-of-range alcohol recovery device through a pipeline; the bottom heavy component outlet of the light component rectifying tower (23) is connected with the lower feed inlet of the ethylene glycol product tower (24) through a pipeline; the top outlet of the ethylene glycol product tower (24) is connected with an extra-terrestrial 1, 2-BDO recovery processing device through a pipeline; the bottom outlet of the ethylene glycol product tower (24) is connected with an external recovery processing device through a pipeline; the upper outlet of the ethylene glycol product tower (24) is an ethylene glycol product outlet; the exhaust gas outlet of the membrane separator (28) is connected with an out-of-range recovery device through a pipeline, and the recovered gas outlet of the membrane separator (28) is connected with the top feed inlet of the hydrogenation reactor (17) through a pipeline.
25. The system of claim 24, wherein the carbonylation reactor (1) is externally connected to a dehydration column (10); the dehydration tower (10) is provided with a feed inlet and a dry gas outlet; the top outlet of the esterification reaction tower (9) and the recovered gas outlet of the pressure swing adsorption tank (16) are connected with the feed inlet of the dehydration tower (10) through pipelines; and a dry gas outlet of the dehydration tower (10) is connected with a feed inlet at the top of the carbonylation reactor (1) through a pipeline.
26. The system of claim 25, wherein the dehydration column comprises two molecular sieve dryers a and B operating and regenerating alternately; the molecular sieve dryer A and the molecular sieve dryer B are filled with adsorbents.
27. The apparatus system for producing dimethyl oxalate through high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol through hydrogenation according to claim 25, wherein the outlet of the bottom of the carbonylation reactor (1) is connected with an outlet heat exchanger I (3); the outlet heat exchanger I (3) is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the CO raw material pipeline and N2Drying of raw material pipes and dewatering towers (10)The dry gas outlet is connected with the cold material flow inlet of the outlet heat exchanger I (3) through a pipeline; a cold flow outlet of the outlet heat exchanger I (3) is connected with a top feed inlet of the carbonylation reactor (1) through a pipeline; a discharge port at the bottom of the carbonylation reactor (1) is connected with a hot material flow inlet of the outlet heat exchanger I (3) through a pipeline; and the hot material outlet of the outlet heat exchanger I (3) is connected with the feed inlet of the first gas-liquid separator (4) through a pipeline.
28. The apparatus system for producing dimethyl oxalate by high-pressure carbonylation of industrial synthesis gas and preparing ethylene glycol by hydrogenation according to claim 27, wherein the carbonylation reactor (1) is externally connected with a steam drum I (2); the steam pocket I (2) is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a steam outlet; a refrigerant inlet of the steam pocket I (2) is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum I (2) is connected with a refrigerant inlet at the bottom of the carbonylation reactor (1) through a pipeline; a top refrigerant outlet of the carbonylation reactor (1) is connected with a vapor-liquid mixture inlet of the steam drum I (2) through a pipeline; and a steam outlet of the steam drum I (2) is connected with an out-of-range steam recovery system through a pipeline.
29. The system of claim 28, wherein a carbonylation recycle compressor (8) is connected between the branched outlet a of the methanol washing column (7) and the lower feed inlet of the esterification reaction column (9); the carbonylation circulating compressor (8) is provided with an inlet and an outlet; the branch outlet A is connected with the inlet of the carbonylation circulating compressor (8) through a pipeline; the outlet of the carbonylation circulating compressor (8) is connected with the lower feed inlet of the esterification reaction tower (9) through a pipeline.
30. The system of claim 29, wherein a compressor (14) is connected between the top outlet of the NO recovery column (13) and the bottom inlet of the MN recovery column (15); the compressor (14) is provided with an inlet and an outlet; the top outlet of the NO recovery tower (13) is connected with the inlet of the compressor (14) through a pipeline; the outlet of the compressor is connected with the bottom feed inlet of the MN recovery tower (15) through a pipeline.
31. The system of claim 30, wherein the outlet of the bottom of the hydrogenation reactor (17) is connected to an outlet heat exchanger ii (20); the outlet heat exchanger II (20) is provided with a cold material flow inlet, a cold material flow outlet, a hot material inlet and a hot material flow outlet; the bottom outlet of the DMO rectifying tower (6), the recovered gas outlet of the membrane separator (28) and the outlet of the hydrogenation circulating compressor (25) are connected with the cold stream inlet of the outlet heat exchanger II (20) through pipelines; a cold flow outlet of the outlet heat exchanger II (20) is connected with a top feed inlet of the hydrogenation reactor (17) through a pipeline; a discharge hole at the bottom of the hydrogenation reactor (17) is connected with a hot material flow inlet of the outlet heat exchanger II (20) through a pipeline; and the hot material flow outlet of the outlet heat exchanger II (20) is connected with the feed inlet of the second gas-liquid separator through a pipeline.
32. The system of claim 31, wherein a start-up heater (19) is connected to a top feed inlet of the hydrogenation reactor (17); the start-up heater (19) is provided with a feed inlet and a discharge outlet; a cold material flow outlet of the outlet heat exchanger II (20) is connected with a feed inlet of the start-up heater (19) through a pipeline; and the discharge hole of the start-up heater is connected with the top feed inlet of the hydrogenation reactor (17) through a pipeline.
33. The system of claim 32, wherein the hydrogenation reactor (17) is externally connected with a steam drum II (18); the steam pocket II (18) is provided with a refrigerant inlet, a refrigerant outlet, a vapor-liquid mixture inlet and a steam outlet; a refrigerant inlet of the steam pocket II (18) is connected with a refrigerant raw material pipeline through a pipeline; a refrigerant outlet of the steam drum II (18) is connected with a refrigerant inlet at the bottom of the hydrogenation reactor (17) through a pipeline; a top refrigerant outlet of the hydrogenation reactor (17) is connected with a vapor-liquid mixture inlet of the steam drum II (18) through a pipeline; and a steam outlet of the steam drum II (18) is connected with an out-of-range steam recovery system through a pipeline.
34. The system of claim 33, wherein the second gas-liquid separator comprises a high-pressure gas-liquid separator (21) and a low-pressure gas-liquid separator (26); the high-pressure gas-liquid separator (21) is provided with a feed inlet, a gas outlet and a liquid outlet; the low-pressure gas-liquid separator (26) is provided with a feed inlet, a gas outlet and a liquid outlet; a discharge hole at the bottom of the hydrogenation reactor (17) is connected with a feed hole of the high-pressure gas-liquid separator (21) through a pipeline; a gas outlet of the high-pressure gas-liquid separator (21) is provided with a branch outlet K and a branch outlet L, the branch outlet K is connected with an inlet of the hydrogenation circulating compressor (25) through a pipeline, and the branch outlet L is connected with a feed inlet of the low-pressure gas-liquid separator (26) through a pipeline; a liquid outlet of the high-pressure gas-liquid separator (21) is connected with a middle feed inlet of the methanol separation tower (22) through a pipeline; the gas outlet of the low-pressure gas-liquid separator (26) is connected with the feed inlet of the membrane separator (28) through a pipeline; and a liquid outlet of the low-pressure gas-liquid separator (26) is connected with a middle feed inlet of the methanol separation tower (22) through a pipeline.
35. The system of claim 31, wherein a methanol absorption tank (27) is arranged in front of the feed inlet of the membrane separator (28); the methanol absorption tank (27) is provided with a feed inlet and a purified gas outlet; a gas outlet of the low-pressure gas-liquid separator (26) and a top noncondensable gas outlet of the methanol separation tower (22) are connected with a feed inlet of the methanol absorption tank (27) through a pipeline; the purified gas outlet of the methanol absorption tank (27) is connected with the feed inlet of the membrane separator (28) through a pipeline.
36. The system of claim 24, wherein the carbonylation reactor (1) is a plate reactor, a tubular reactor or a composite tubular-plate reactor.
37. The system of claim 36, wherein the carbonylation reactor (1) is a plate fixed bed carbonylation reactor; the center of the plate type fixed bed carbonylation reactor is provided with a plate group fixing cavity, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed carbonylation reactor; the catalyst bed layer is filled with a carbonylation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate type fixed bed carbonylation reactor, a bottom refrigerant inlet of the plate type fixed bed carbonylation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate type fixed bed carbonylation reactor through a pipeline; and the top inlet of the plate fixed bed carbonylation reactor is connected with the top inlet of the catalyst bed layer through a pipeline, and the top outlet of the plate group fixing cavity is connected with the top refrigerant outlet of the plate fixed bed carbonylation reactor through a pipeline.
38. The system of claim 24, wherein the esterification tower (9) is a packed tower or a mixed tower of tray and packing having both tray and packing sections.
39. The system of claim 24, wherein the methanol washing column (7), the methanol rectifying column (5), the methanol recovery column (11), the NO recovery column (13), the MN recovery column (15), the DMO rectifying column (6) and the nitric acid concentrating column (12) are packed columns, plate columns or bubble cap columns.
40. The system of claim 24, wherein the hydrogenation reactor (17) is a plate-type bed reactor, a tubular reactor or a plate-tubular combined reactor.
41. The system of claim 35, wherein the hydrogenation reactor (17) is a plate-type fixed bed hydrogenation reactor; a plate group fixing cavity is arranged in the center of the plate type fixed bed hydrogenation reactor, a plate group is arranged in the plate group fixing cavity, and the plate group fixing cavity is also provided with a bottom inlet and a top outlet; a catalyst bed layer is arranged between the outer wall of the plate group fixing cavity and the inner wall of the plate type fixed bed hydrogenation reactor; the catalyst bed layer is filled with a hydrogenation reaction catalyst and is also provided with a top inlet and a bottom outlet; at the bottom of the plate-type fixed bed hydrogenation reactor, a bottom refrigerant inlet of the plate-type fixed bed hydrogenation reactor is connected with a bottom inlet of the plate group fixed cavity through a pipeline, and a bottom outlet of the catalyst bed layer is connected with a bottom discharge hole of the plate-type fixed bed hydrogenation reactor through a pipeline; the top of plate fixed bed hydrogenation ware, plate fixed bed hydrogenation ware's top feed inlet with the top entry pipe line connection of catalyst bed layer, the top export in the fixed chamber of plate group with plate fixed bed hydrogenation ware's top refrigerant export pipe line connection.
42. The system of claim 24, wherein the membrane separator (28) comprises 1-100 hollow fiber membrane modules connected in parallel or in series.
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| CN204211669U (en) | 2015-03-18 |
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