CN103113187B - By the method for acetic acid ethanol co-production ethyl acetate - Google Patents
By the method for acetic acid ethanol co-production ethyl acetate Download PDFInfo
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- CN103113187B CN103113187B CN201310053290.1A CN201310053290A CN103113187B CN 103113187 B CN103113187 B CN 103113187B CN 201310053290 A CN201310053290 A CN 201310053290A CN 103113187 B CN103113187 B CN 103113187B
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Abstract
The present invention relates to a kind of method of being produced ethanol co-production ethyl acetate by acetic acid direct hydrogenation, described method comprises: in reactor, pass into acetic acid feed and hydrogen, to carry out hydrogenation reaction in the presence of a catalyst, gained reaction product is isolated thick product through flash tank, this thick product obtains ethanol and ethyl acetate product through purification, wherein gasify with the described acetic acid feed counter current contact be atomized by spray nozzle after preheating from the circulating hydrogen on described flash tank top, and then mix with high concentration hydrogen and form gaseous state mixing raw material, and described gaseous state mixing raw material carries out preheating before entering described reactor.Described method makes full use of stream stock heat and carries out heat exchange optimization, reduce energy consumption and save facility investment, and hydrogen can be obtained through pressure-variable adsorption by the periodic off-gases of coke-oven gas methanation, solves hydrogen source problem.Can be realized by acetic acid One-step production ethanol co-production ethyl acetate by method of the present invention.
Description
Technical field
The present invention relates to the method that acetic acid direct hydrogenation produces ethanol co-production ethyl acetate.
Background technology
1, background of invention
In recent years, oil price continues soaring, and environmental stress is day by day heavy, and the development in pluralism of the energy is all being sought in countries in the world, develops new fungible energy source, as wind energy, sun power and biomass energy etc.With fastest developing speed in biomass energy is alcohol fuel, and it is one of emerging reproducible green energy resource, can be used as gasoline dope and automotive fuel, has the advantage that octane value is high and anti-knocking property are good.At present, existing more than 20 countries promote and use ethanol petrol in the world.The production of existing industrialization alcohol fuel is raw material mainly with corn, wheat and sucrose greatly, although technical maturity, ethanol production is subject to the restriction of raw material, and production cost is higher, is difficult to satisfied long-term demand for development; In the long run, with Mierocrystalline cellulose (comprising agricultural crop straw, forestry processing waste material etc.) for raw material production alcohol fuel, may be solve one of raw material sources and the approach carrying out large-scale production.But the production efficiency of Mierocrystalline cellulose ethanol technology is low at present, and production cost is higher, only carries out technological innovation, the space just made further progress.
In order to alleviate oil shortage and environmental stress, China is using the important component part of the new forms of energy of efficient for exploitation, low consumption and safety as energy resource safety strategy.From 2003, promote the use of the strategic measure that vehicle-use alcohol gasoline just becomes China.Biomass ferment method and ethylene hydration method prepare ethanol traditional method, the former take grain as the production line of raw material, the latter take oil as the production line of raw material, in view of the importance of China's grain security and the scarcity of petroleum resources, these two production lines do not meet the requirement of China's Sustainable development.
China's oil natural resources shortage, be net import of oil state, coal resources are but very abundant, its total resources accounts for 90% of fossil energy, therefore, play the advantage of coal resources, saving and alternatives oil, develops coal-based clean energy, realizes the coordinated development of economy, environment and the energy.At present, industrial producing synthesis gas from coal methanol, acetic acid is prepared in carbonylation of methanol, and these two kinds of explained hereafter are with low cost, and technology maturation, carries out large-scale industrial production.The process operation of the direct shortening ethanol synthesis of acetic acid is simple, and raw material sources are extensive, cheap, so the direct catalytic hydrogenation process of acetic acid is the feasible path of ethanol production, is also realize by one of effective way of the indirect preparing liquid fuel additive of coal.Study the production technique of novel acetic acid catalysis preparation of ethanol by hydrogenating, significant.
2, prior art route and problem
2.1, biomass ferment method
Biomass material is grow up on the basis of wine brewing through the method for production of ethanol from microbial fermentation, and within the quite long period of history, biomass ferment was once the unique commercial run producing ethanol.Biomass material is divided into three classes: (1) carbohydrate, comprises sugarcane, sweet sorghum, beet etc.; (2) starch based, comprises corn, wheat, sweet potato, cassava etc.; (3) cellulose family, comprises stalk, crudefiber crop, farm crop shell skin, branch fallen leaves etc.
At present, domestic biomass fuel ethanol is fixed a point to produce by Jilin fuel limited liability company, Henan Tian Guan alcohol fuel company, four enterprises of Fengyuan Biological Chemistry Co., Ltd., Anhui Prov. and Heilongjiang China Resources Alcohol Co., Ltd, standard production 1,020,000 tons/year of alcohol fuels.Wherein, Henan Tian Guan alcohol fuel company is mainly raw material with wheat, and other three company is all raw material with corn.These enterprises all facilitate production and the application of internal fuel ethanol.
Due to traditional grain fermentation method technical obsolescence, slag and effluent process is more difficult, and the productive rate of wood-based composites and selectivity are not high, cost and the energy consumption of producing ethanol are all very high, be difficult to ensure the fast-developing demand to the energy of national economy, therefore the exploitation of biomass energy are dry biomass, i.e. unedible biomass, instead of food crop.As a supplement or adjust, particularly as the recycling means of waste, biomass fuel and bio-ethanol are all feasible, but they all can not as the main body of future source of energy, to development bio-ethanol or biofuel, take Mierocrystalline cellulose as raw material, directly acquisition ethanol or polyvalent alcohol are only feasible path, can build the new forms of energy platform replacing grain fermentation alcohol fuel.
Over nearly 30 years, be that the report of raw material production alcohol fuel is a lot of with cellulose substances, have more deep research to the aspect such as method for hydrolysis, zymotechnique.Celunol company of the U.S. take transgenic escherichia coli as fermented bacterium, adopts secondary dilute sulphuric acid hydrolysis process to carry out the production of alcohol fuel.In November, 2006, the said firm builds the demo plant of a biomass-making ethanol in Louisiana, produces alcohol fuel fifty-five million gallon per year.2000, it was raw material that Canadian Iogen company sets up one with wheat straw, take dilute sulphuric acid as catalyzer, and the demo plant of enzyme hydrolysis process produces alcohol fuel, produces 300-400 ten thousand liters of ethanol per year.In the last few years, China also achieved greater advance in Mierocrystalline cellulose ethanol.Set up based on dilute acid hydrolysis process in Spring in Pig House of Shanghai Fengxian district, produce the demo plant of 600 tons of alcohol fuels per year, realize continuous prodution, for the mass-producing of China's Mierocrystalline cellulose preparing fuel ethanol is laid a good foundation.
At present, be that the technology of raw material production ethanol is mature on the whole with Mierocrystalline cellulose, but the maximum drawback of this technique is that cellulase is expensive, consumption is many, cause the price of cellulosic ethanol cannot be that ethanol cost prepared by raw material is competed mutually with grain.Therefore, need cheaper effective Mierocrystalline cellulose pretreatment process, Optimization Technology link, strengthens the Social benefit and economic benefit of this technology.
2.2, synthetic gas biological process
The gordian technique of synthetic gas bio-fermentation process is optimize gasification process and select suitable fermentation strain.At present, bio-engineering corporation of the U.S. (BRI) has developed synthetic gas fermentative production of ethanol technology, successfully can produce ethanol rapidly with cellulose castoff, viable economically, and has environmental protection effect.In November, 2003, the said firm set up the demonstration project of first biomass synthesis gas fermentation ethanol in the Arkansas State, and started to establish first commercial operations device the end of the year 2005.Such as, but still there are some problems needing to solve in synthetic gas zymotechnique, how to increase CO and H
2productive rate, improve the economy of alcohol production; How to use genetically engineered to develop better bacterium, and optimize anaerobic fermentation process by medium optimization, cell cycle and immobilized cell technology etc.; How to develop the reactor of high mass transfer efficiency, utilize osmotic evaporation film or extractive fermentation new technology to eliminate Product inhibiton, improve alcohol yied.
2.3, chemosynthesis legal system ethanol
Chemical synthesis is divided into ethylene hydration method and synthetic gas synthesis method, and the former comprises again ethene indirect hydration method and ethene direct hydration method.
2.3.1, ethylene hydration method
Olefin hydration method prepares the important method of alcohols material, is applied to Germany's acetylene hydration in 1913 the earliest and prepares acetaldehyde in organic chemical industry.Nineteen thirty, U.S. combinating carbide company established the full scale plant that first employing oil boom cracked product ethene is raw material ethanol, i.e. ethene indirect hydration method.Ethene indirect hydration method is also known as sulfuric acid process, and reaction is carried out in two steps, first passes in the vitriol oil by ethene under certain temperature, pressure, form sulfuric ester, then in hydrolysis tower, sulfuric ester heating hydrolysis is obtained ethanol, this technological reaction mild condition, conversion of ethylene is high.Because ethene indirect hydration ethanol is more old, can produce a large amount of dilute sulphuric acids in process of production, etching apparatus, cause environmental pollution, Production Flow Chart is long, is now replaced by ethene direct hydration method.
2.3.2, synthetic gas direct synthesis technique
Be that to prepare ethanol be the novel process of rising in the last few years to raw material with synthetic gas.The technique of the direct synthesizing alcohol of synthetic gas is also by one of path of the indirect preparing liquid fuel of coal.This technique is with CO, CO
2and H
2for unstripped gas.Synthetic gas direct hydrogenation ethanol is strong heat release and is easier to the reaction carried out.But CO easily methanation reaction occurs in hydrogenation process, consume a large amount of H
2, selectivity and the productive rate of ethanol be improved, must select to suppress the catalyzer of methanation reaction and suitable reaction conditions.
From aerodynamic point, in reaction process, C-C key synthesis speed is slow, C
2the rate of chain growth of intermediate is fast, causes synthetic gas reactant CO, CO on a catalyst
2transformation efficiency and the selectivity of ethanol all very low, reaction product is generally mixture of lower alcohols.Use different catalyzer in reaction process, the selectivity difference of Methanol product and ethanol is very large.Therefore select suitable carrier, auxiliary agent, simple efficient, that ethanol selectivity is high, the life-span the is longer catalyzer that research and development are applicable to this technique has very large development prospect.
2.4, ester catalysis Hydrogenation alcohol
In the product of the direct shortening of carboxylic acid in addition to alcohols, part by product ester is also comprised.In order to improve the selectivity of object product alcohol, obtain more fuel alcohol, people often adopt catalyst hydrogenation reduction Ester.From the 1950's, people have started the large-scale commercial production from natural acid ester fatty alcohol, wherein can produce some basic chemical productss, as methyl alcohol and ethanol etc. by ester through hydrogenation reaction.The product of ester through hydrogenation reaction may comprise alcohol, acid and hydrocarbon polymer etc., and product composition depends primarily on the structure of matter, catalyzer and reaction conditions.The reaction that ester through hydrogenation is reduced to alcohol is the process of the C-O key selectivity fracture that carbonyl is adjacent.
In ester through hydrogenation reduction reaction, adopt copper-based catalysts can obtain the alcohol product of highly selective.Copper-based catalysts at first for the hydrogenation reaction of methyl-formiate, after this widespread use in the industrial production.Prepare in the process of fatty alcohol, often adopt copper chromite catalyst, this catalyst system is by cupric oxide CuO and cupric chromate CuCr
2o
4deng mole to mix, but the material wherein with ester through hydrogenation activity is CuO.The research such as Evans shows, at Catalysts Cu/SiO
2or on Lei Nitong, ester through hydrogenation reactive behavior is also very high, and be conducive to protection of the environment.
The nineties in 20th century, people propose some newer production paths to ethanol preparation technology.Kiff etc. are with acetic acid and alkene for raw material first carries out esterification, and the ester that then will obtain hydrogenation on agent CuO/ZnO catalyzer generates ethanol and another kind of higher alcohols, and higher alcohols dewater formation alkene again, and olefin recycle uses.
Bradley etc. propose similar technical process with Juran etc., maximum different being of this technique and above-mentioned technical process first forms methyl acetate by low-carbon alcohol such as methyl alcohol and acetic acid generation esterification, then methyl acetate hydrogenating reduction production object product ethanol and by-product carbinol, methyl alcohol recirculation is used for esterification.
2.5, domestic and international progress
Hydrogenation of carboxylic acids reduction reaction is the very important reaction of a class in organic chemistry, has purposes widely in the industry.From 20 century 70s, start the document and the patent report that occur producing alcohol about the direct shortening of carboxylic acid abroad.Researchist adopts different catalyzer, affects Carboxylic Acid and leads the factor with object selectivity of product, and propose corresponding reaction mechanism in the direct catalytic hydrogenation reaction of exploration carboxylic acid.Acetic acid is the simplest carboxylic acid, and wide material sources are cheap, and therefore people often select acetic acid as research object, investigates different catalysts, differential responses condition on the impact of acetic acid conversion and selectivity of product, and sets up reaction kinetics equation.At present, people have developed various catalyzer, and can be corresponding alcohol by aromatic carboxylic acid and fatty acid reduction, sour transformation efficiency and selectivity of product be all higher.Due to the restriction of reaction conditions and catalyzer cost, the technique of acid catalysis reduction alcohol processed is also only limitted to the laboratory study stage.
Schreck etc. adopt multistep processes by acetic acid catalysis preparation of ethanol through hydrogenation.Principal reaction step is: (1), in tubular reactor, under the effect of an acidic catalyst perfluorinated sulfonic resin, acetic acid and 2-methyl-2-butene react and generate ester; (2) it is 200 DEG C in temperature, pressure is under the reaction conditions of 0.69MPa, liquid-phase catalyst 65wt%CuO/35wt%ZnO is after 190 DEG C of reduction 6h, and the ester in hydrogenating reduction (1), obtains ethanol and another kind of alcohol (identical with the carbon number of alkene in 1); (3) distillation method separating alcohol and high-boiling point alcohol is adopted; (4) form catalyzer by equimolar melting boric acid and phosphoric acid, dewatered by high boiling alcohol at the temperature of 100-200 DEG C, 2-methyl-2-butene of getting back, reaction cycle is carried out.
Ryashentseva etc. have studied rhenium sulfide Re
2s
7on the impact of acetic acid hydrogenation reaction.Catalyzer is Powdered, and preparation method has two kinds: ammonium perrhenate, potassium perrhenate or rehenic acid and hydrogen sulfide generation coprecipitation reaction; Ammonium perrhenate solution is immersed on alumina catalyst support, is obtained by reacting after drying at 120 DEG C with hydrogen sulfide.Reduction reaction is carried out in open glass ampoule, and glass ampoule is placed in the steel autoclave of vibration.Adopt acid base neutralization titration to obtain residual acetic acid content, product is measured by on-line chromatographic analysis.Acetic acid feed amount is 4g, catalyst levels 0.13g, is 230 DEG C in temperature, and hydrogen pressure is that after reacting 6h under the condition of 13MPa, ethanol maximum output is 88.0%.
The method that Celanese International Corp's acetic acid hydrogenation obtains crude ethanol product comprises, and in the presence of a catalyst the acetic acid hydrogenation in acetic acid feed stream is formed crude ethanol product; At least part of crude ethanol product is separated into the first overhead product comprising ethanol, water and ethyl acetate in the first tower, and comprises the first resistates of acetic acid; In the second tower, at least part of first overhead product is separated into the second overhead product comprising ethyl acetate and the second resistates comprising second alcohol and water; And at least part of second resistates is separated into the 3rd overhead product comprising ethanol and the 3rd resistates comprising water in the 3rd tower.Product is that content is up to the ethanol of 96wt% and acetic acid conversion is low, does not substantially relate to the problem that heat is optimized.
Summary of the invention
For above-mentioned prior art situation, the object of the present invention is to provide a kind of by the method for single stage method realization by acetic acid hydrogenation ethanol co-production ethyl acetate.
On the one hand, the invention provides a kind of method of being produced ethanol co-production ethyl acetate by acetic acid direct hydrogenation, described method comprises: in reactor, pass into acetic acid feed and hydrogen, to carry out hydrogenation reaction in the presence of a catalyst, the reaction product obtained is isolated thick product through flash tank, described thick product obtains ethanol and ethyl acetate through purification, wherein gasify with the described acetic acid feed counter current contact be atomized by spray nozzle after preheating from the circulating hydrogen on described flash tank top, and then mix with high concentration hydrogen and form gaseous state mixing raw material, and described gaseous state mixing raw material carries out preheating before entering described reactor.
In a preferred embodiment, described circulating hydrogen carries out preheating by the described reaction product heat exchange with described reactor outlet, then contacts with the described acetic acid feed of atomization.
In a preferred embodiment, described circulating hydrogen by with the described reaction product heat exchange of described reactor outlet after, then carry out heat exchange with other interchanger, then contact with the described acetic acid feed of atomization.
In a preferred embodiment, described gaseous state mixing raw material carries out by the described reaction product heat exchange with described reactor outlet the temperature being preheated to 80-300 DEG C before entering the reactor.
In a preferred embodiment, described high concentration hydrogen obtains by being separated from the periodic off-gases of coke-oven gas methanation.
In a preferred embodiment, described high concentration hydrogen is obtained by the periodic off-gases separation from described coke-oven gas methanation of membrane sepn, pressure-variable adsorption or molecular sieve adsorption.
In a preferred embodiment, the concentration being separated from the periodic off-gases of described coke-oven gas methanation the high concentration hydrogen obtained is more than 95%.
In a preferred embodiment, described acetic acid feed is the mixture comprising acetic acid and water, and wherein said acetic acid feed comprises the acetic acid of at least 70wt%.
In a preferred embodiment, in described gaseous state mixing raw material, the mol ratio of hydrogen and acetic acid is 100: 1-1: 1.
In a preferred embodiment, the raw materials components mole ratio of described hydrogen and acetic acid raw material is 20: 1-5: 1.
In a preferred embodiment, the temperature of reaction in described reactor is 100-350 DEG C, and reaction pressure is 0.1-3MPa, gas air speed 1,000-20,000h
-1.
In a preferred embodiment, the catalyzer used is made up of carrier and metal composites, and described metal composites is selected from Pt/Sn, Pt/Ru, Pt/Re, Pd/Ru, Pd/Re, Co/Pd, Co/Pt, Co/Cr, Co/Pu, Ag/Pd, Cu/Pd, Ni/Pd, Au/Pd, Ru/Re, Ru/Fe or Pt/Sn/Ni.
In a preferred embodiment, described metal composites is Pt/Sn/Ni.
In a preferred embodiment, based on the weight of catalyzer, the carrier of Ni and 80-98.8wt% of Sn, 0.1-5wt% of Pt, 0.01-5wt% of consisting of of described catalyzer: 0.01-10wt%.
In a preferred embodiment, described carrier is silicon oxide, aluminum oxide, sial complex carrier, ferric oxide, titanium oxide, zirconium white, magnesium oxide, graphite, gac, graphitized charcoal or their mixture.
In a preferred embodiment, described carrier is sial complex carrier.
In a preferred embodiment, the mean diameter of described sial complex carrier is 3-8mm, and density is 0.42-0.80g/ml, and water-absorbent is 0.40-0.95g water/g carrier, and specific surface area is 150-280m
2/ g and pore volume are 0.40-0.80ml/g.
In a preferred embodiment, the described carrier modifier be selected from alkaline earth metal oxide, alkalimetal oxide, alkali earth metasilicate, alkali metal silicate, metal oxide or their mixture carries out modification.
In a preferred embodiment, described modifier is Ca (NO
3)
2, CaSiO
3, Mg (NO
3)
2, MgSiO
3or their mixture.
In a preferred embodiment, by one or more rectifying tower, purification is carried out to described thick product.
In a preferred embodiment, described purification is completed by four rectifying tower, wherein:
In the first rectifying tower, become to comprise the first cut of ethyl acetate and acetaldehyde by described thick product separation, and comprise the first tower bottoms of second alcohol and water;
In Second distillation column, described first fraction seperation is become to comprise the second cut of acetaldehyde and comprises the second tower bottoms of ethyl acetate;
In the 3rd rectifying tower under extraction agent exists, described first tower bottoms is separated into the 3rd cut comprising ethanol, and comprises the 3rd tower bottoms of water and extraction agent; With
Described 3rd tower bottoms is separated into the 4th cut comprising water and the 4th tower bottoms comprising extraction agent in the 4th rectifying tower.
In a preferred embodiment, described second cut comprises the acetaldehyde of 0-85wt%, and described second tower bottoms comprises the ethyl acetate of 80-100wt%.
In a preferred embodiment, described 3rd cut comprises the ethanol of 90-99.8wt%, and water content is not higher than 0.8wt%.
In a preferred embodiment, being circulated back at least partially in described reactor of described first cut and described second cut.
In a preferred embodiment, being circulated back at least partially in described 3rd rectifying tower of described 4th tower bottoms.
Technical process good stability of the present invention, flow process is short, energy consumption is low, the temperature head making full use of various flows stock carries out substep heat exchange, to realize the target of efficiency utilization heat, and hydrogen can be isolated to by the periodic off-gases of coke(oven)gas methane, solves hydrogen source problem.In addition, ethanol yield is high and production cost is low, can save acid esterification device compared with two-step approach, because acetic acid conversion is greater than 99%, can reduce separating device, reduce investment outlay when carrying out separating-purifying to product.And the catalyzer of reaction can be recycled, and according to the output of market requirement flexible two kinds of products, the device ability to ward off risks can be increased.In addition, be increased to more than 95% by the periodic off-gases of coke(oven)gas methane through pressure-variable adsorption concentration for the raw hydrogen of reacting, then react, solve hydrogen source problem.
Accompanying drawing explanation
Fig. 1 is according to the process flow sheet of the present invention for acetic acid hydrogenation ethanol co-production ethyl acetate, wherein, and 101-reactor; 102,103,104,105,108-interchanger; 106-flash tank; 107-recycle compressor; 109-spray nozzle; 110-pressure-variable adsorption workshop section; 201-first rectifying tower; 202-Second distillation column; 203-the 3rd rectifying tower; 204-the 4th rectifying tower.
Embodiment
The present invention adopts single stage method by acetic acid hydrogenation ethanol co-production ethyl acetate, by changing operational condition, can control the content of ethyl acetate in product, thus reaching the object of coproduction ethyl acetate.In coarse ethanol product, acetic acid content is lower than 1wt%, and carries out purification to this thick product further and finally obtain proper fuel alcohol product and ethyl acetate product.Acetic acid conversion of the present invention is greater than 99%, and when major product is ethanol, its selectivity is greater than 95%.
Method by acetic acid direct hydrogenation production ethanol co-production ethyl acetate of the present invention comprises: in reactor, pass into acetic acid feed and hydrogen, to carry out hydrogenation reaction in the presence of a catalyst, the reaction product obtained is isolated thick product through flash tank, last described thick product obtains ethanol and ethyl acetate product through purification, wherein gasify with the described acetic acid feed counter current contact be atomized by spray nozzle after preheating from the circulating hydrogen on described flash tank top, and then mix with high concentration hydrogen and form gaseous state mixing raw material, and described gaseous state mixing raw material carries out preheating before entering described reactor.
According to a specific embodiment of the present invention, the technical process of acetic acid direct hydrogenation ethanol co-production ethyl acetate is as shown in Figure 1, specific as follows:
The coarse ethanol product that acetic acid raw material and reactor 101 export enters spray nozzle 109 after interchanger 104 heat exchange, circulate from flash tank 106 circulating hydrogen out through compressor 107, cycling hot hydrogen stream stock is formed through interchanger 103 and optional interchanger 108 with coarse ethanol product, the acetic acid that spray nozzle 109 is atomized and cycling hot hydrogen counter current contact gasifying, and then form gaseous state mixing raw material with the high concentration hydrogen obtained by coke-oven gas methanation periodic off-gases centrifugal station 110, the coarse ethanol product heat utilizing reactor 101 to export, heat further through interchanger 102 pairs of gaseous state mixing raw materials, then reactor 101 is entered, and acetic acid hydrogenation is carried out under catalyst action, form the thick product comprising ethanol.Acetic acid raw material can be selected from acetic acid, diacetyl oxide, acetaldehyde, ethyl acetate or their mixture, also directly can enter synthesizing alcohol unit from methanol carbonylation unit (generation raw acetic acid).Preferably, acetic acid raw material is reacted after can mixing with rare gas element, reacts preferred gas phase and carries out; Preferably, temperature of reaction 100-350 DEG C, reaction pressure 0.1-3MPa, gas air speed 1,000-20,000h
-1, the hydrogen liquor ratio of raw hydrogen and acetic acid is 100: 1-1: 1, is preferably greater than 20: 1-5: 1.Acetic acid vaporization temperature, lower than the boiling point under working pressure, when just having started to react, heats hydrogen and acetic acid by interchanger 108 and the acetic acid well heater (not shown) that goes into operation respectively, to reach the required feeding temperature of reaction.
As shown in Figure 1, coarse ethanol product is out rear preferably first through flash tank 106 from reactor, and this flash tank 106 temperature is preferably 100-350 DEG C, and pressure is 100-2000KPa, and the temperature of preferred flash tank is identical with reactor 101 with pressure.Comprise hydrogen and a small amount of hydro carbons from flash tank 106 gas out, loop back reactor 101 by compressor 107.From flash tank 106 liquid phase out and reactor 101 export form similar, directly can enter the first rectifying tower 201 (i.e. lightness-removing column) and carry out separating-purifying, also first can pass into relay tank (not shown) to re-refine, be mainly containing the product of ethanol, water and possible little amount unreacted acetic acid from the first rectifying tower 201 tower bottoms out.
The overhead fraction being separated the first rectifying tower 201 is mainly the light constituent such as ethyl acetate, acetaldehyde, and because acetaldehyde and ethyl acetate are intermediate products, therefore acetaldehyde and ethyl acetate can return reaction process.Also as required the first cut of the overhead extraction from the first rectifying tower 201 can be passed into Second distillation column 202 (i.e. dealdehyder tower), can obtain ethyl acetate product at the bottom of dealdehyder tower 202 tower, the acetaldehyde product of its overhead extraction can return reaction workshop section.
The first tower bottoms containing alcohol product enters the 3rd rectifying tower 203 (i.e. dehydration tower), and extraction agent adds from the middle and upper part of dehydration tower 203, under extraction agent effect, obtains proper fuel alcohol product from the tower top of the 3rd rectifying tower 203.
At the bottom of the tower of dehydration tower 203, extraction contains the extraction agent of water and a small amount of ethanol, the 4th rectifying tower 204 (i.e. extraction agent recovery tower) is entered through pump (not shown), from its overhead extraction water, preferably, the water obtained enters water treatment operation, and from this tower middle and lower part extraction extraction agent, preferably looped back dehydration tower 203 and reuse.
In the present invention, reactor 101 can be fixed-bed reactor, fluidized-bed reactor, shell and tube-type reactor or ebullated bed reactor, and they can be adiabatic or heat exchange type reactor.In the present invention, reactor 101 is preferably fixed-bed reactor.
In the present invention, the catalyzer used in reactor 101 is preferably made up of carrier and metal composites, and described metal composites is selected from Pt/Sn, Pt/Ru, Pt/Re, Pd/Ru, Pd/Re, Co/Pd, Co/Pt, Co/Cr, Co/Pu, Ag/Pd, Cu/Pd, Ni/Pd, Au/Pd, Ru/Re, Ru/Fe or Pt/Sn/Ni.Wherein preferred described metal composites is Pt/Sn/Ni, and more preferably, based on the weight of whole catalyzer, the content of Pt, Sn, Ni in described catalyzer and described carrier is respectively 0.01-10wt%, 0.01-5wt%, 0.1-5wt% and 80-98.8wt%.
In the present invention, the solid support material for catalyzer can comprise: silicon oxide, aluminum oxide, sial complex carrier, ferric oxide, titanium oxide, zirconium white, magnesium oxide, graphite, gac, graphitized charcoal or their mixture.Support of the catalyst can use modifier modification, and described modifier can be selected from alkaline earth metal oxide, alkalimetal oxide, alkali earth metasilicate, alkali metal silicate, metal oxide or their mixture etc., preferably Ca (NO
3)
2, CaSiO
3, Mg (NO
3)
2, MgSiO
3or their mixture.Described carrier is preferably sial complex carrier, and more preferably, its mean diameter is 3-8mm, and density is 0.42-0.80g/ml, and water-absorbent is 0.40-0.95g water/g carrier, and specific surface area is 150-280m
2/ g, pore volume is 0.40-0.80ml/g.
In the present invention, the rectifying tower of use can be tray column or packing tower, and column plate can be sieve plate, fixing valve plate and float valve plate, and stage number is generally 20-85, and packing tower can be structured packing or dumped packing.Column plate and filler can be distributed in a tower or different towers, and the heat of the reboiler in tower can from technological process or outside, and technological process can increase more reactor, flash tank, condenser, reboiler etc. according to actual needs.Except specified otherwise, tower service temperature is the temperature of liquid in reboiler, and the pressure of top of tower is normal atmosphere.
When first rectifying tower 201 operates at normal atmospheric pressure, entering tower liquid temp is 105-115 DEG C, and the temperature of tower top first cut is 70-95 DEG C, and the first cut enters Second distillation column 202.When Second distillation column 202 under atmospheric pressure operates, the second tower bottoms temperature is 70-110 DEG C, and tower top second cut temperature is 60-100 DEG C, and working pressure is 1-375KPa.Because acetaldehyde and ethyl acetate are intermediate products, therefore acetaldehyde or ethyl acetate can return reaction process.Also can take out ethyl acetate product as required at the bottom of tower, the acetaldehyde product of overhead extraction can return reaction process, and tower top and ethyl acetate content in tower reactor are than being preferably less than 0.1: 1.
First still liquid enters dehydration tower 203 middle and lower part after taking out, reflux ratio is preferably 1: 10-10: 1, under atmospheric pressure operate, overhead fraction temperature is 75-95 DEG C, tower bottoms temperature is 80-105 DEG C, and dehydration tower 203 can be used as extraction tower, adds extraction agent (such as ethylene glycol, water etc.), extraction agent can be recycled, and dehydration column overhead obtains proper fuel alcohol product.Dehydration tower 203 also can be board-like or packing tower, and stage number is such as 20-70.
Extraction at the bottom of dehydration tower 203 tower contains the extraction agent of water and a small amount of ethanol, enters extraction agent recovery tower 204, overhead extraction water through pump.If containing a small amount of acetic acid in the first tower bottoms, then this small amount of acetic acid is from the overhead extraction of extraction agent recovery tower 204, goes to water treatment operation together, from tower middle and lower part extraction extraction agent, loops back dehydration tower 203 and reuses.
Alternatively, these rectifying tower all can introduce side line, thus removing impurity, improve product purity.
The present invention passes through under the above-described reaction conditions by acetic acid direct hydrogenation ethanol, and acetic acid raw material is reacted after can mixing with rare gas element, reacts preferred gas phase and carries out.The transformation efficiency of acetic acid is higher than 99%, and later separation operation is corresponding simpler, thus saves facility investment.Can also according to the market requirement, regulate the indexs such as temperature of reaction, pressure and hydrogen liquor ratio, improve the selectivity of ethyl acetate thus reach the object of coproduction ethyl acetate.And the present invention makes full use of process system internal heat energy, the optimum combination of various flows stock heat is made to carry out heat exchange, simple to operate, stable reaction.
Below by embodiment, the present invention is described in further detail, but should be appreciated that these embodiments only for illustrating the present invention, and be not used in and limit the scope of the invention.
Embodiment 1
After gasifying containing the charging of 98.6wt% acetic acid, 1.2wt% water and surplus acetaldehyde and the circulation gas of 120 DEG C, be pass into insulation fix bed reactor together with the hydrogen of 96% (v/v) with coke-oven gas methanation periodic off-gases through the concentration that pressure-variable adsorption obtains, under catalyst action, reaction generates the coarse ethanol product comprising ethanol, acetic acid, water and ethyl acetate, react under the temperature of reaction of 220 DEG C and 1.8MPa pressure, described catalyzer comprises 0.5wt% platinum, 0.4wt% tin and the 1.2wt% nickel of load on the calcium metasilicate modification sial complex carrier of 6mm.Reactor inlet is got back in unreacted hydrogen recirculation, makes at 4021h
-1gHSV (i.e. gas air speed) under the mol ratio of total hydrogen and acetic acid be 5.6: 1.Under these conditions, the acetic acid of 99.1% transforms, and now the selectivity of ethanol is 75.2%, and the selectivity of ethyl acetate is 23.3%, and the selectivity of acetaldehyde is 1.2%.
Use technical process separating-purifying coarse ethanol product as shown in Figure 1.Described coarse ethanol product is passed in the first rectifying tower with the feeding rate of 10.5kg/min.This first tower is packing tower, the temperature of 84 DEG C in this tower of air pressing operation.
First cut is carried out condensing reflux at the first top of tower with 5: 1 ratios, is passed in Second distillation column with the feeding rate of 3.2kg/min by part cut, this second tower is packing tower, operates under higher than atmospheric 192KPa.Second tower bottoms is taken out with the speed of 2.8kg/min.Second cut refluxes, and another part then turns back in reactor, and wherein 99.8wt% is acetaldehyde, and all the other are ethyl acetate.
It is solid mass-transferring sieve-tray tower that first tower bottoms is passed into the 3rd rectifying tower the 203, three tower with the feeding rate of 7.0kg/min, at 101 DEG C in this tower of air pressing operation.Add extraction agent (as ethylene glycol) to improve alcohol product content, the 3rd tower bottoms is taken out with the speed of 2.9kg/min.3rd cut takes out after condensing reflux is carried out at top.
4th rectifying tower is packing tower, operates under large pressure the temperature of 105 DEG C.Take out after 4th cut carries out condensing reflux and deliver to wastewater treatment operation, the 4th tower bottoms takes out Posterior circle and returns the 3rd tower.Table 1 is the analytical results table of the charging of rectifying tower, cut and tower bottoms composition, and nd material is with "-" display (other minute quantity material is unlisted).
Table 1
As can be seen from the result of table 1, on the one hand, by method of the present invention, acetic acid conversion is high, the acetic acid content of the first tower bottoms is very low simultaneously, and the ethanol content that the 3rd tower cut obtains is 98.1%, and water-content is lower than 0.8%, meet process of alcohol products and require (GB18350-2001), and the quality of alcohol product and ethyl acetate is much better than the product obtained by existing method.On the other hand, according to the thermograde of raw material, make full use of the heat of the rear coarse ethanol product of reaction, substep preheating is carried out to raw material, effectively utilizes heat, reach energy-saving and cost-reducing target.In addition, because acetic acid conversion is high, depickling rectifying tower can be saved, thus save facility investment.
Embodiment 2
After gasifying containing the charging of 99.2wt% acetic acid, 0.6wt% water and surplus acetaldehyde and the circulation gas of 120 DEG C, insulation fix bed reactor is passed into together with the hydrogen that coke-oven gas methanation periodic off-gases concentration after pressure-variable adsorption is 96% (v/v), under catalyst action, reaction generates the coarse ethanol product comprising ethanol, acetic acid, water and ethyl acetate, react under the temperature of reaction of 285 DEG C and 2.0MPa pressure, described catalyzer comprises the 0.5wt% platinum of load on the calcium metasilicate modification sial complex carrier of 6mm and 0.4wt% tin.Reactor inlet is got back in unreacted hydrogen recirculation, makes at 9036h
-1gHSV under the mol ratio of total hydrogen and acetic acid be 8.2: 1.Under these conditions, the acetic acid of 99.3% transforms, and now the selectivity of ethanol is 92.9%, and the selectivity of ethyl acetate is 5.2%, and the selectivity of acetaldehyde is 0.4%.According to technical process separating-purifying coarse ethanol product as shown in Figure 1.
Described coarse ethanol product is passed in the first rectifying tower with the feeding rate of 12.2kg/min.First tower is packing tower, the temperature of 84 DEG C in this tower of air pressing operation.First cut is carried out condensing reflux at the first top of tower, and is passed in Second distillation column with the feeding rate of 0.5kg/min by part cut, the second tower is packing tower, operates under higher than atmospheric 165KPa.Second tower bottoms is taken out with the speed of 0.46kg/min.Second cut refluxes, and another part then turns back in reactor.
First tower bottoms is passed into the 3rd rectifying tower with the feeding rate of 11.5kg/min, and the 3rd tower is solid mass-transferring sieve-tray tower, at 102 DEG C in this tower of air pressing operation.Add extraction agent (as ethylene glycol etc.) to improve alcohol product content, the 3rd tower bottoms is taken out with the speed of 3.8kg/min.3rd cut takes out after condensing reflux is carried out at top.
4th rectifying tower is packing tower, operates under large pressure the temperature of 105 DEG C.Take out after 4th cut carries out condensing reflux and deliver to wastewater treatment operation, the 4th tower bottoms takes out Posterior circle and returns the 3rd tower.Table 2 is the analytical results table of the charging of rectifying tower, cut and tower bottoms composition, and nd material is with "-" display (other minute quantity material is unlisted).
Table 2
As can be seen from the result of table 2, on the one hand, by method of the present invention, acetic acid conversion is high, the acetic acid content of the first tower bottoms is very low simultaneously, and the ethanol content that the 3rd tower cut obtains is 99.0%, and water-content is not higher than 0.8%, meet process of alcohol products and require (GB18350-2001), and the quality of alcohol product and ethyl acetate is much better than the product obtained by existing method.On the other hand, according to the thermograde of raw material, make full use of the heat of the rear coarse ethanol product of reaction, substep preheating is carried out to raw material, effectively utilizes heat, reach energy-saving and cost-reducing target.In addition, be better than acetic acid conversion high, depickling rectifying tower can be saved, thus save facility investment.
In addition, can find out according to embodiment 1 and embodiment 2, although the platinum tin two metal active constituent catalyzer used in embodiment 2 is different from three metal active constituent catalyzer in embodiment 1, but the purity of final ethanol and ethyl acetate product is suitable, namely, in the methods of the invention, two metal active constituent catalyzer have good catalytic performance equally.
Embodiment 3
After gasifying containing the charging of 97.5wt% acetic acid, 2.0wt% water and surplus acetaldehyde and the circulation gas of 118 DEG C, insulation fix bed reactor is passed into together with the hydrogen that coke-oven gas methanation periodic off-gases concentration after pressure-variable adsorption is 97%, under following catalyst action, reaction generates the coarse ethanol product comprising ethanol, acetic acid, water and ethyl acetate, react under the temperature of reaction of 295 DEG C and 2.1MPa pressure, described catalyzer comprises 0.7wt% platinum, 0.6wt% tin and the 1.5wt% nickel of load on the calcium metasilicate modification sial complex carrier of 6mm.Reactor inlet is got back in unreacted hydrogen recirculation, makes at 16220h
-1gHSV under the mol ratio of total hydrogen and acetic acid be 14.5: 1.Under these conditions, the acetic acid of 99.6% transforms, and now the selectivity of ethanol is 94.7%, and the selectivity of ethyl acetate is 3.5%, and the selectivity of acetaldehyde is 0.7%.Use technical process separating-purifying coarse ethanol product as shown in Figure 1.
Described coarse ethanol product is passed in the first rectifying tower with the feeding rate of 8.1kg/min.First tower is packing tower, the temperature of 82 DEG C in this tower of air pressing operation.First cut is carried out condensing reflux at the first top of tower, and is passed in Second distillation column with the feeding rate of 0.32kg/min by part cut, the second tower is packing tower, operates under higher than atmospheric 150KPa.Second tower bottoms is taken out with the speed of 0.25kg/min.A second cut part refluxes, and another part then turns back in reactor.
First tower bottoms is passed into the 3rd rectifying tower with the feeding rate of 7.52kg/min, and the 3rd tower is solid mass-transferring sieve-tray tower, at 101 DEG C in this tower of air pressing operation.Add extraction agent (as ethylene glycol etc.) to improve alcohol product content, the 3rd tower bottoms is taken out with the speed of 2.25kg/min.3rd cut takes out after condensing reflux is carried out at top.
4th rectifying tower is packing tower, operates under large pressure the temperature of 104 DEG C.Take out after 4th cut carries out condensing reflux and deliver to wastewater treatment operation, the 4th tower bottoms takes out Posterior circle and returns the 3rd tower.Table 3 is the analytical results table of the charging of rectifying tower, cut and tower bottoms composition, and nd material is with "-" display (other minute quantity material is unlisted).
Table 3
From the result of table 3 and the result of table 1 similar, wherein the ethanol content that obtains of the 3rd tower cut is up to 99.4%, and water-content, not higher than 0.5%, meets process of alcohol products and requires (GB18350-2001).And compare from the result of table 3 and the result of table 1, when in known catalyzer, platinum content improves, the selectivity of acetic acid conversion and ethanol increases all to some extent.
Embodiment 4
After gasifying containing the charging of 99.4wt% acetic acid, 0.5wt% water and surplus acetaldehyde and the circulation gas of 125 DEG C, insulation fix bed reactor is passed into together with the hydrogen that coke-oven gas methanation periodic off-gases concentration after pressure-variable adsorption is 98%, under following catalyst action, reaction generates the coarse ethanol product comprising ethanol, acetic acid, water and ethyl acetate, react under the temperature of reaction of 292 DEG C and 2.1MPa pressure, described catalyzer comprises 0.7wt% platinum, 0.6wt% tin and the 1.5wt% nickel of load on the calcium metasilicate modification sial complex carrier of 6mm.Reactor inlet is got back in unreacted hydrogen recirculation, makes at 2597h
-1gHSV under the mol ratio of total hydrogen and acetic acid be 18.4: 1.With this understanding, the acetic acid of 100% transforms, and now the selectivity of ethanol is 96.4%, and the selectivity of ethyl acetate is 2.5%, and the selectivity of acetaldehyde is 0.6%.Use technical process separating-purifying coarse ethanol product as shown in Figure 1.
Described coarse ethanol product is passed in the first rectifying tower with the feeding rate of 8.2kg/min.Provide the composition of liquid feeding in table 2.First tower is packing tower, the temperature of 82 DEG C in this tower of air pressing operation.First cut is carried out condensing reflux at the first top of tower, and is passed in the second tower with the feeding rate of 0.29kg/min by part cut, the second tower is packing tower, operates under higher than atmospheric 167KPa.Second tower bottoms is taken out with the speed of 0.23kg/min.A second cut part refluxes, and another part then turns back in reactor.
First tower bottoms is passed into the 3rd tower with the feeding rate of 7.78kg/min, and the 3rd tower is solid mass-transferring sieve-tray tower, at 101 DEG C in this tower of air pressing operation.Add extraction agent (as ethylene glycol etc.) to improve alcohol product content, the 3rd tower bottoms is taken out with the speed of 2.30kg/min.3rd cut takes out after condensing reflux is carried out at top.
4th tower is packing tower, operates under large pressure the temperature of 104 DEG C.Take out after 4th cut carries out condensing reflux and deliver to wastewater treatment operation, the 4th tower bottoms takes out Posterior circle and returns the 3rd tower.Table 4 is the analytical results table of the charging of rectifying tower, cut and tower bottoms composition, and nd material is with "-" display (other minute quantity material is unlisted).
Table 4
From the result of table 4 and the result of table 1 similar, wherein the ethanol content that obtains of the 3rd tower cut is up to 99.6%, and water-content, not higher than 0.4%, meets process of alcohol products and requires (GB18350-2001).And carry out contrast from the result of table 4 and the result of table 3, after total air speed minimizing of reaction and the mol ratio of hydrogen and acetic acid improve, acetic acid conversion can reach 100%, and the selectivity of ethanol improves, and the selectivity of ethyl acetate and acetaldehyde reduces simultaneously.
Other embodiments
(1) except the mol ratio of hydrogen total in charging and acetic acid is changed into 20: 1,50: 1,80: 1 and 95: 1 respectively, to implement the inventive method with embodiment 1 same program, the result of acquisition is all similar to Example 1.
(2) except being changed into respectively except Pt/Ru, Pd/Re, Co/Cr, Co/Pu, Cu/Pd, Ni/Pd, Ru/Re and Ru/Fe by metal active constituent, to implement the inventive method with embodiment 2 same program, the result of acquisition is all similar to Example 2.
(3) except the carrier of catalyzer is changed into magnesium oxide, gac, silicon oxide respectively, and outside the mixture of aluminum oxide and graphitized charcoal, to implement the inventive method with embodiment 1 same program, the result of acquisition is all similar to Example 1.
(4) except changing into except diacetyl oxide by raw acetic acid, to implement the inventive method with embodiment 1 same program, the result of acquisition is all similar to Example 1.
From above-described embodiment, when the mol ratio of total hydrogen and acetic acid increases, the ethanol selectivity of reaction improves, and the selectivity of ethyl acetate reduces, and the mol ratio can passing through hydrogen and acetic acid in feed change regulates the output of ethanol and ethyl acetate.Air speed change does not cause obvious impact to reaction, and can be suitable in a big way, turndown ratio is large.In alcohol product, ethanol content is greater than 98wt%, and water-content, lower than 0.8wt%, meets the requirement of GB18350-2001 to alcohol fuel.Ethyl acetate product meets industrial requirements simultaneously.
In addition, the present invention is by forming cycling hot hydrogen stream stock after the coarse ethanol product heat exchange of circulating hydrogen and reactor outlet, the acetic acid be atomized by spray nozzle and cycling hot hydrogen counter current contact also gasify, and then be separated with the periodic off-gases of coke-oven gas methanation the high concentration hydrogen obtained and form gaseous state mixing raw material, before entering reactor, utilize coarse ethanol product heat to heat raw material further, then enter reactor and react; Acetic acid conversion of the present invention and ethanol selectivity high, make full use of stream stock heat and carry out heat exchange optimization, energy consumption is low and save facility investment, and hydrogen can be isolated to by the periodic off-gases of coke(oven)gas methane, solves hydrogen source problem.
It should be pointed out that the professional and technical personnel for making the art, under the prerequisite not departing from the technology of the present invention principle, is what can realize the multiple amendment of these embodiments, and these amendments also should be considered as in the scope that the present invention should protect.
Claims (23)
1. produced the method for ethanol co-production ethyl acetate by acetic acid direct hydrogenation for one kind, described method comprises: in reactor, pass into acetic acid feed and hydrogen, to carry out hydrogenation reaction in the presence of a catalyst, the reaction product obtained is isolated thick product through flash tank, described thick product obtains ethanol and ethyl acetate through purification
Wherein gasify with the described acetic acid feed counter current contact be atomized by spray nozzle after preheating from the circulating hydrogen on described flash tank top, and then mix with high concentration hydrogen and form gaseous state mixing raw material, and described gaseous state mixing raw material carries out preheating before entering described reactor
Described circulating hydrogen carries out preheating by the described reaction product heat exchange with described reactor outlet, then contacts with the described acetic acid feed of atomization, and described high concentration hydrogen obtains by being separated from the periodic off-gases of coke-oven gas methanation.
2. method according to claim 1, is characterized in that, described circulating hydrogen by with the described reaction product heat exchange of described reactor outlet after, then carry out heat exchange with other interchanger, then contact with the described acetic acid feed of atomization.
3. method according to claim 1, is characterized in that, described gaseous state mixing raw material carries out by the described reaction product heat exchange with described reactor outlet the temperature being preheated to 80-300 DEG C before entering the reactor.
4. method according to claim 1, is characterized in that, described high concentration hydrogen is obtained by the periodic off-gases separation from described coke-oven gas methanation of membrane sepn, pressure-variable adsorption or molecular sieve adsorption.
5. the method according to claim 1 or 4, is characterized in that, the concentration being separated the high concentration hydrogen obtained from the periodic off-gases of described coke-oven gas methanation is more than 95%.
6. method according to claim 1, is characterized in that, described acetic acid feed is the mixture comprising acetic acid and water, and wherein said acetic acid feed comprises the acetic acid of at least 70wt%.
7. method according to claim 1, is characterized in that, in described gaseous state mixing raw material, the mol ratio of hydrogen and acetic acid is 100:1-1:1.
8. method according to claim 7, is characterized in that, the raw materials components mole ratio of described hydrogen and acetic acid raw material is 20:1-5:1.
9. method according to claim 1, is characterized in that, the temperature of reaction in described reactor is 100-350 DEG C, and reaction pressure is 0.1-3MPa, gas air speed 1,000-20,000h
-1.
10. method according to claim 1, it is characterized in that, the catalyzer used is made up of carrier and metal composites, and described metal composites is selected from Pt/Sn, Pt/Ru, Pt/Re, Pd/Ru, Pd/Re, Co/Pd, Co/Pt, Co/Cr, Co/Pu, Ag/Pd, Cu/Pd, Ni/Pd, Au/Pd, Ru/Re, Ru/Fe or Pt/Sn/Ni.
11. methods according to claim 10, is characterized in that, described metal composites is Pt/Sn/Ni.
12. methods according to claim 10, is characterized in that, based on the weight of catalyzer, and the carrier of Ni and 80-98.8wt% of Sn, 0.1-5wt% of Pt, 0.01-5wt% of consisting of of described catalyzer: 0.01-10wt%.
13. methods according to claim 10, is characterized in that, described carrier is silicon oxide, aluminum oxide, sial complex carrier, ferric oxide, titanium oxide, zirconium white, magnesium oxide, graphite, gac, graphitized charcoal or their mixture.
14. methods according to claim 13, is characterized in that, described carrier is sial complex carrier.
15. methods according to claim 14, is characterized in that, the mean diameter of described sial complex carrier is 3-8mm, and density is 0.42-0.80g/ml, and water-absorbent is 0.40-0.95g water/g carrier, and specific surface area is 150-280m
2/ g and pore volume are 0.40-0.80ml/g.
16. methods according to claim 13, it is characterized in that, the described carrier modifier be selected from alkaline earth metal oxide, alkalimetal oxide, alkali earth metasilicate, alkali metal silicate, metal oxide or their mixture carries out modification.
17. methods according to claim 16, is characterized in that, described modifier is Ca (NO
3)
2, CaSiO
3, Mg (NO
3)
2, MgSiO
3or their mixture.
18. methods according to claim 1, is characterized in that, carry out purification by one or more rectifying tower to described thick product.
19. methods according to claim 18, is characterized in that, described purification is completed by four rectifying tower, wherein:
In the first rectifying tower, become to comprise the first cut of ethyl acetate and acetaldehyde by described thick product separation, and comprise the first tower bottoms of second alcohol and water;
In Second distillation column, described first fraction seperation is become to comprise the second cut of acetaldehyde and comprises the second tower bottoms of ethyl acetate;
In the 3rd rectifying tower under extraction agent exists, described first tower bottoms is separated into the 3rd cut comprising ethanol, and comprises the 3rd tower bottoms of water and extraction agent; With
Described 3rd tower bottoms is separated into the 4th cut comprising water and the 4th tower bottoms comprising extraction agent in the 4th rectifying tower.
20. methods according to claim 19, it is characterized in that, described second cut comprises the acetaldehyde of 0-85wt%, described second tower bottoms comprises the ethyl acetate of 80-100wt%.
21. methods according to claim 19, is characterized in that, described 3rd cut comprises the ethanol of 90-99.8wt%, and water content is not higher than 0.8wt%.
22. methods according to claim 19, is characterized in that, being circulated back at least partially in described reactor of described first cut and described second cut.
23. methods according to claim 19, is characterized in that, being circulated back at least partially in described 3rd rectifying tower of described 4th tower bottoms.
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| CN105566063A (en) * | 2014-10-14 | 2016-05-11 | 中国石油化工股份有限公司 | Ethanol preparation method |
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