EP1515935A1 - Procede pour produire du 1,4-butanediol par hydrogenation combinee en phase gazeuse et en phase liquide - Google Patents

Procede pour produire du 1,4-butanediol par hydrogenation combinee en phase gazeuse et en phase liquide

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Publication number
EP1515935A1
EP1515935A1 EP03740211A EP03740211A EP1515935A1 EP 1515935 A1 EP1515935 A1 EP 1515935A1 EP 03740211 A EP03740211 A EP 03740211A EP 03740211 A EP03740211 A EP 03740211A EP 1515935 A1 EP1515935 A1 EP 1515935A1
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EP
European Patent Office
Prior art keywords
hydrogenation
reactor
catalyst
stage
gbl
Prior art date
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Application number
EP03740211A
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German (de)
English (en)
Inventor
Michael Hesse
Stephan Schlitter
Holger Borchert
Markus Schubert
Markus Rösch
Nils Bottike
Rolf-Hartmuth Fischer
Alexander Weck
Gunther Windecker
Gunnar Heydrich
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BASF SE
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BASF SE
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Publication of EP1515935A1 publication Critical patent/EP1515935A1/fr
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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
    • 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
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/18—Polyhydroxylic acyclic alcohols
    • C07C31/20—Dihydroxylic alcohols

Definitions

  • the present invention relates to a two-stage process for the preparation of optionally alkyl-substituted butanediol by catalytic hydrogenation of substrates which are selected from the group consisting of derivatives of maleic acid and succinic acid and these acids themselves.
  • the first hydrogenation stage is in the gas phase, the second hydrogenation stage in the Liquid phase carried out.
  • derivatives are understood to mean anhydrides which, like the acids, can have one or more alkyl substituents.
  • EP-A 0 322 140 discloses a continuous process for the preparation of tetrahydrofuran (THF) and for the simultaneous production of THF and GBL by gas phase hydrogenation of MSA and BSA.
  • the claimed catalyst contains copper, zinc and aluminum and a further element from group DA, IDA, VA, VIH, IHK to VUB, the lanthanum and actinium series, or Ag and Au.
  • these catalyst systems can be used to obtain 90-95% THF yields based on pure MA, mixtures of GBL and THF can be obtained at a pressure of approx. 20 bar.
  • EP-A 0 404 408 discloses a catalyst for the MSA hydrogenation system, the catalytically active material of which essentially corresponds to the material of US 5,072,009. It is used fixed on a support as a coated catalyst. Only chromium-containing catalysts are used in the exemplary embodiments. High GBL yields can be achieved at a pressure of 2 bar. If a higher pressure is used, the THF yield increases and the GBL yield decreases.
  • No. 5,149,836 discloses a multi-stage gas phase process for the production of GBL and THF with variable product selectivities, a mixture of pure MAA and hydrogen being passed over a catalyst which contains copper, zinc and aluminum in a first stage. This crude discharge is then passed over a chromium-containing catalyst to produce the THF.
  • WO 99/38856 discloses a catalyst consisting only of copper and chromium, with which GBL selectivities of 92 to 96 mol% are obtained in a straight pass, starting from pure MSA.
  • EP-A 638 565 discloses a catalyst containing copper, chromium and silicon, the composition according to one example corresponding to approximately 78% CuO, 20% Cr 2 O 3 and 2% SiO 2 . With pure MSA and nitrogen-hydrogen mixtures, a GBL yield of 98% could be obtained.
  • GB-A 1 168 220 discloses a gas phase process for the preparation of GBL in which MSA or BSA are hydrogenated to GBL over a binary copper-zinc catalyst. In all of the exemplary embodiments, work is carried out at atmospheric pressure, with GBL yields of 94 mol% starting from pure MSA.
  • DE-OS 2 404 493 also discloses a process for the preparation of GBL by catalytic hydrogenation of mixtures of MA, BSA, maleic acid (MS), succinic acid (BS) and water over a metallic catalyst, in addition to copper-chromite catalysts copper-zinc and copper-zinc-aluminum precipitation catalysts are also used.
  • WO 91/16132 discloses the hydrogenation of MA to GBL, using a catalyst containing CuO, ZnO and Al 2 O 3 , which is reduced at 150 ° C. to 350 ° C. and activated at 400 ° C. Activation is intended to increase the service life of the catalyst system.
  • a CuO-ZnO-containing catalyst is disclosed in US 6,297,389. It can be used to convert pure MSA to GBL after activation, with GBL yields of 92 to 96% being obtained in a single pass based on pure MSA.
  • WO 95/22539 discloses a process for the production of GBL by catalytic hydrogenation of MSA and / or BSA on a catalyst made of copper, zinc and Zirconium exists. Based on pure MSA, GBL yields of up to 99% are achieved.
  • WO 99/35136 discloses a two-stage process for the preparation of GBL and THF, in which MSA is hydrogenated in a first stage with a copper-containing catalyst and this reaction product is passed over an acidic silicon-aluminum-containing catalyst.
  • WO 97/24346 describes a copper oxide-aluminum oxide catalyst with which the hydrogenation of MA can be achieved in yields of 92 mol% to GBL.
  • DE-A 1 277 233 discloses a process for the preparation of mixtures of various alcohols by hydrogenating lactones with hydrogen. Copper chromite mixed with barium on an inactive aluminum oxide support is used as the catalyst.
  • GB-A 1 230 276 describes a process for producing BDO from GBL over a copper oxide-chromium oxide catalyst at a temperature between 180 ° C. and 230 ° C.
  • the service life of the catalysts is to be extended with copper chromite catalysts doped with potassium, sodium, rubidium, aluminum, titanium, iron, cobalt or nickel.
  • a mixture of dioxane, GBL, water and carboxylic acids is used to produce BDO.
  • the reaction described takes place at high pressure (170 bar) in the liquid phase, preferably using the solvent dioxane, and copper-chromium oxide catalysts are also used.
  • Copper chromite catalysts are doped with Pd according to J01121-228-A in order to achieve a higher conversion. Further copper chromite catalysts are described by Dasunin, Maeva, Z.Org.chim. 1 (1965), No. 6, S996-1000; YES 5366/69; YES 7240770; J4 9024-906; J49087-610; in the exemplary embodiments, pure GBL is converted to BDO in the liquid phase.
  • a catalyst consisting of CuO and ZnO is described in WO 82/03854. With this, a BDO selectivity of 98.4% can be achieved in the gas phase at a pressure of 28.5 bar and a temperature of 217 ° C. The turnover of pure GBL is unsatisfactorily low.
  • Copper impregnation catalysts doped with palladium and potassium are described in documents US Pat. No. 4,797,382; US 4,885,411 and EP-A 0 318 129. They are suitable for converting GBL to butanediol.
  • JP-A 0 634 567 describes a copper-iron-aluminum-containing catalyst which is suitable for the hydrogenation of pure GBL to BDO at high pressure (250 bar).
  • a process for the preparation of BDO starting from maleic acid esters is listed in WO 99/35113. Here, hydrogenation is carried out in three successive stages. Starting from maleic acid ester, succinic acid ester is produced on a noble metal-containing catalyst, which is converted in a second stage to GBL and THF. GBL is separated off and converted to BDO in a third stage at higher pressure.
  • WO 99/35114 describes a process for the preparation of BDO by liquid phase hydrogenation of GBL, succinic acid ester or mixtures of the two at pressures between 60 bar and 100 bar and temperatures between 180 ° C. and 250 ° C.
  • a copper oxide-zinc oxide catalyst is used as the catalyst.
  • EP-A 0 382 050 works on the hydrogenation of pure GBL over a catalyst containing cobalt oxide, copper oxide, manganese oxide and molybdenum oxide.
  • DE 2 845 905 describes a continuous process for the preparation of butanediol starting from maleic anhydride.
  • MSA is dissolved in monohydric aliphatic alcohols with hydrogen at pressures of 250 bar and 350 bar on copper chromite catalysts.
  • a process for the simultaneous production of BDO and THF starting from MSA on catalysts containing copper, chromium and manganese is disclosed in EP-A 0 373 947. Mixtures of MSA and GBL, mixtures of MSA and 1,4-dioxane and pure MSA are used. Mixtures of THF and BDO are obtained in all cases. The high yields of tetrahydrofuran are disadvantageous in this process.
  • Chromium-containing catalysts are disclosed in documents CN-A 1 113 831, CN-A 1 116 615, CN-A 1 138 018 and CN-A 1 047 328.
  • CN-A 1 137 944 uses a copper, chromium, manganese, barium and titanium-containing catalyst.
  • a copper, chromium, zinc and titanium-containing catalyst can be used for the hydrogenation of mixtures of GBL and MSA.
  • CN-A 1 182 732 describes a process for the preparation of BDO by gas phase hydrogenation of MA with copper and chromium-containing catalysts at 200 to 250 ° C. and a pressure of 30 to 70 bar.
  • MSA is dissolved in a suitable solvent during the hydrogenation.
  • DE-A 2 455 617 describes a three-stage process for the preparation of BDO.
  • a first stage solutions of MSA in GBL to BSA in GBL are hydrogenated on a Ni-containing catalyst.
  • this solution of BSA in GBL is hydrogenated in the liquid phase to GBL, then water, succinic anhydride and succinic acid are separated from the GBL and the pure GBL is partially recycled and in a third Process stage on a copper-zinc oxide catalyst in the liquid phase at high pressure converted to butanediol.
  • No. 4,301,077 uses a ruthenium-containing catalyst for the hydrogenation of MA to BDO.
  • DE-A 3 726 510 discloses the use of a copper, cobalt and phosphorus-containing catalyst for the direct hydrogenation of MA.
  • a pure copper oxide-zinc oxide catalyst is used with J0 2025-434-A.
  • pure MSA can be implemented at a pressure of 40 bar.
  • the yield of butanediol is only 53.5 mol%, and 40.2 mol% GBL are found as a secondary yield.
  • EP-A 373 946 discloses a process in which a rhenium-doped copper oxide-zinc oxide catalyst MSA of the gas phase is converted directly to BDO.
  • J0 2233-627-A (here a copper-zinc-aluminum catalyst is used)
  • J0 2233- 630-A here is a manganese, barium and silicon-containing copper -Chrome catalyst used
  • J0 2233-631-A (here a copper and aluminum containing catalyst is used).
  • a catalyst containing copper, manganese and potassium is described in J0-A 2233-632.
  • EP-A 431 923 describes a two-stage process for the preparation of BDO and THF, GBL being formed in a first stage by liquid-phase hydrogenation of MA and this being converted to butanediol in a second stage by gas-phase reaction on a copper-silicon-containing catalyst becomes.
  • US 5,196,602 discloses a process for the preparation of butanediol by hydrogenation of MA or maleic acid with hydrogen in a two-stage process.
  • MSA is hydrogenated to BSA and / or GBL, which in a second stage is converted to BDO in the presence of an Ru-containing catalyst.
  • MAA which has been prepurified as a starting material for the hydrogenation reactions and which, after its preparation, has generally been freed of impurities by distillation.
  • MSA is produced by partial oxidation of certain hydrocarbons, namely benzene, butene mixtures and n-butane, the latter being preferably used.
  • the crude product of the oxidation contains above all by-products such as water, carbon monoxide, carbon dioxide, unreacted Starting hydrocarbon as well as acetic and acrylic acid, these by-products being independent of the hydrocarbons used in the oxidation.
  • the by-products are mostly separated off by complex processes, for example by distillation, as mentioned above.
  • Promoted copper catalysts are used as catalysts, as described in Journal of Catalysis 150, pages 177 to 185 (1994). These are chromium-containing catalysts of the types Cu / Mn / Ba / Cr and Cu / Zn Mg / Cr. Thus, according to the disclosure of this application, chromium-containing catalysts are used for the hydrogenation of MA qualities which have the impurities set out above. However, the use of chromium-containing catalysts is avoided as much as possible due to the toxicity.
  • chromium-containing catalysts Due to their toxicity, newer technologies are increasingly moving away from the use of chromium-containing catalysts. Examples of chromium-free catalyst systems can be found in the publications WO 99/35139 (Cu-Zn-oxide), WO 95/22539 (Cu-Zn-Zr) and US Pat. No. 5,122,495 (Cu-Zn-Al oxide).
  • WO 99/35139 Cu-Zn-oxide
  • Cu-Zn-Zr Cu-Zn-Zr
  • US Pat. No. 5,122,495 Cu-Zn-Al oxide
  • BDO was obtained by direct hydrogenation of pure GBL - which in turn was obtained by hydrogenation of MSA and subsequent elaborate purification.
  • pure MSA was used as the starting material, which contained only small amounts of impurities, since otherwise no satisfactory selectivity and catalyst life were achieved.
  • Chromium-containing catalysts were used, in particular also in the second stage, in order to achieve high BDO selectivity and the desired service life. To avoid the use of chromium-containing catalysts, it was alternatively possible to use noble metal-containing catalysts which are comparable in terms of yield, selectivity and stability to chromium-containing catalysts, but are significantly more cost-intensive.
  • the process should not contain any chromium Catalysts, preferably also do not require catalysts which contain noble metals and have a high selectivity for BDO, in particular provide little THF.
  • a catalyst is used in both hydrogenation stages, the ⁇ 95 wt .-%, preferably 5 to 95 wt .-%, in particular 10 to 80 wt .-% CuO and> 5 wt .-%, preferably 5 to 95 wt. -%, in particular 20 to 90 wt .-% of a carrier, there is a higher pressure in the second reactor than in the first
  • Reactor and the product mixture taken from the first reactor is introduced into the second reactor without further purification.
  • hydrogen is separated off before performing step c). This is when the product stream is transferred to the Liquid phase particularly easy.
  • the hydrogen is advantageously used again in stage a).
  • water is separated off before stage c) is carried out. In this way, longer catalyst service lives and also better catalyst performance, such as yield and selectivity, can generally be achieved.
  • the hydrogenation process according to the invention can comprise an upstream stage which comprises the preparation of the starting material to be hydrogenated by partial oxidation of a suitable hydrocarbon and the separation of the starting material to be hydrogenated from the product stream obtained therewith.
  • a coarse separation is preferably carried out, which does not require an unnecessarily high outlay and in which an amount of impurities which is not tolerable in the processes known hitherto remains in the starting material.
  • This educt to be hydrogenated is in particular MSA.
  • MSA which originates from the partial oxidation of hydrocarbons, is preferably used. Suitable hydrocarbon streams are benzene, C-olefins (eg n-butenes, C 4 raffinate streams) or n-butane. N-Butane is used with particular preference since it is an inexpensive, economical starting material. Processes for the partial oxidation of n-butane are for example in Ulimann's Encyclopedia of Industrial Chemistry, 6 Edition, Electronic Release, Maleic and Fumaric Acids - Maleic Anhydride.
  • reaction product thus obtained is then preferably taken up in a suitable organic solvent or mixture which has a boiling point at least 30 ° C. higher than MAA at atmospheric pressure.
  • This solvent is brought to a temperature in the range between 20 and 160 ° C, preferably between 30 and 80 ° C.
  • the gas stream from the partial oxidation containing maleic anhydride can be brought into contact with the solvent in a variety of ways: (i) introducing the gas stream into the solvent (for example via gas inlet nozzles or gassing rings), (ii) spraying the solvent into the gas stream and (iii) countercurrent contact between the upward gas stream and the downward solvent in a tray or packing column.
  • the gas absorption apparatus known to those skilled in the art can be used in all three variants. When choosing the solvent to be used, care must be taken to ensure that this does not react with the starting material, for example the preferably used MA.
  • Suitable solvents are: tricresyl phosphate, dibutyl maleate, butyl maleate, high molecular waxes, aromatic hydrocarbons with a molecular weight between 150 and 400 and a boiling point above 140 ° C, such as dibenzylbenzene; Alkyl phthalates and dialkyl phthalates with C 1 -C 18 -alkyl groups, for example dimethyl phthalate, diethyl phthalate, dibutyl phthalate, di-n-propyl and di-isopropyl phthalate, undecyl phthalate, diundecyl phthalate, methyl phthalate, ethyl phthalate, butyl phthalate, n-propyl and isopropyl phthalate; Di-C ⁇ -C 4 alkyl esters of other aromatic and aliphatic dicarboxylic acids, for example dimethyl-2,3-naphthalene-dicarboxylic acid, dimethyl-1, 4-cyclo
  • the solution resulting after treatment with the absorbent generally has an MS A content of about 5 to 400 grams per liter.
  • the waste gas stream remaining after treatment with the absorbent mainly contains the by-products of the previous partial oxidation, such as water, carbon monoxide, carbon dioxide, unreacted butanes, acetic and acrylic acid.
  • the exhaust gas flow is practically free of MSA.
  • the dissolved MSA is then expelled from the absorbent. This is done with hydrogen at or at most 10% above the pressure of the subsequent hydrogenation or alternatively in a vacuum with subsequent condensation of remaining MA.
  • a temperature profile is observed in the stripping column, which results from the boiling points of MSA at the top and the almost MSA-free absorbent at the bottom of the column at the respective column pressure and the dilution set with carrier gas (in the first case with hydrogen).
  • rectification internals can be located above the supply of the raw MSA stream.
  • the almost MSA-free absorbent drawn off from the sump is returned to the absorption zone.
  • the Hj / MSA ratio is about 20 to 400.
  • the condensed MSA is pumped into an evaporator and evaporated there into the circulating gas stream.
  • the MSA hydrogen stream also contains by-products, which are formed during the partial oxidation of n-butane, butenes or benzene with gases containing oxygen, as well as non-separated absorbent. These are primarily acetic acid and acrylic acid as by-products, water, maleic acid and the dialkyl phthalates preferably used as absorbents.
  • the MSA contains acetic acid in amounts of 0.01 to 1% by weight, preferably 0.1 to 0.8% by weight and acrylic acid in amounts of 0.01 to 1% by weight, preferably 0.1 to 0 , 8 wt .-%, based on MSA. In the hydrogenation stage, all or some of the acetic acid and acrylic acid are hydrogenated to ethanol or propanol.
  • the maleic acid content is 0.01 to 1% by weight, in particular 0.05 to 0.3% by weight, based on MA.
  • dialkyl phthalates are used as absorbents, their content in the MA depends heavily on the correct operation of the stripping column, in particular on the rectifying section. Phthalate contents of up to 1.0% by weight, in particular up to 0.5% by weight, should not be exceeded with a suitable mode of operation, since otherwise the consumption of absorbent becomes too high.
  • the hydrogen / maleic anhydride stream preferably obtained as described above is then fed to the first hydrogenation zone and hydrogenated.
  • the catalyst activities and downtimes are practically unchanged compared to the use of strongly pre-cleaned MA, for example by distillation.
  • the gas stream emerging from the first reactor can be processed further in several ways according to the invention.
  • the product stream is condensed out and the liquid stream obtained, together with the excess hydrogen and possibly recycled GBL, is fed to the second hydrogenation, which takes place in the liquid phase.
  • the hydrogen is separated from the second stream before the liquid stream is introduced.
  • the products of the first hydrogenation stage a) are preferably condensed at temperatures of 10 to 60 ° C.
  • the water of reaction formed is largely removed from the condensed product stream before entering the second hydrogenation reactor. This can be done by measures known to the person skilled in the art, for example by distillation of the product stream. Water is preferably removed by condensation of the gas stream emerging from the first reactor. The gas stream is preferably brought to temperatures of 5 to 20 ° C., in particular 10 ° C., below the dew point of GBL at the set pressure. The product stream is partially condensed. BDO, GBL and BSA as well as any higher-boiling by-products and part of the THF are separated from water that remains in the gas phase. The remaining gas stream can be cooled further, for example to separate products such as THF.
  • the hot product stream of the first separation is fed into the liquid phase hydrogenation c). This means that the entire product stream does not have to be cooled down first and then warmed up again.
  • any remaining gas flow is withdrawn from the separator and, in a preferred embodiment, fed to the cycle gas compressor.
  • a Partial stream fed back to the reactor preferably via a cycle gas compressor.
  • the condensed reaction products are continuously removed from the system and fed to the second hydrogenation circuit.
  • the reaction products are then passed over the second catalyst bed.
  • the product stream emerging from the second reactor is separated into liquid stream and gas stream.
  • the gas stream is drawn off and, in a preferred embodiment, fed to the cycle gas compressor.
  • a partial stream is fed back to the reactor, preferably via a cycle gas compressor.
  • By-products formed in the recycle gas stream can be removed by measures known to those skilled in the art, preferably by discharging a small amount of gas.
  • the reaction products contained in the liquid stream are preferably taken continuously from the system and fed to a workup.
  • the by-products found in the condensed liquid phase are mainly GBL, THF, n-butanol in addition to small amounts of propanol.
  • GBL can be returned to the second stage.
  • the inventive method can be operated inexpensively, high BDO yields and selectivities are achieved with low recycle currents.
  • the GBL-laden gas stream from the first stage is compressed to the pressure of the second stage and the recycle gas from the second stage is expanded into the inlet of the first stage, possibly with labor.
  • the gas stream emerging from the second reactor can be cooled, preferably to 10 to 60 ° C.
  • the reaction products are condensed out and passed into a separator.
  • the uncondensed gas stream can be drawn off from the separator and fed compressed to the cycle gas.
  • By-products formed in the recycle gas stream can be removed by measures known to those skilled in the art, preferably by discharging a small amount of gas.
  • the condensed reaction products are removed from the system and processed.
  • the by-products found in the condensed liquid phase are mainly THF and n-butanol in addition to small amounts of propanol.
  • the by-products and water are then separated off from the liquid hydrogenation discharge from the second stage and the desired product BDO is isolated. This happens generally by fractional distillation.
  • By-products and intermediates such as GBL and Di-BDO can be returned to the hydrogenation stage of the first and / or second stage, preferably the second stage, or alternatively worked up by distillation.
  • the process according to the invention can be carried out batchwise, semi-continuously or continuously. Continuous implementation is preferred.
  • this is preferably achieved by a sufficiently high temperature of the starting materials when entering the first hydrogenation reactor.
  • This so-called initial hydrogenation temperature is from 200 to 300 ° C, preferably 235 to 270 ° C.
  • the reaction should preferably be carried out in such a way that a suitably high reaction temperature prevails on the catalyst bed on which the actual reaction takes place.
  • This so-called hot spot temperature is set after the starting materials have entered the reactor and is preferably at values of 210 to 310 ° C., in particular 245 to 280 ° C. The process is preferably carried out such that the inlet temperature and the outlet temperature of the reaction gases are below this hot spot temperature.
  • the hot spot temperature is advantageously in the first half of the reactor, in particular if a tube bundle reactor is present.
  • the hot spot temperature is preferably 5 to 30 ° C., in particular 5 to 15 ° C., particularly preferably 5 to 10 ° C., above the inlet temperature. If the hydrogenation is carried out below the minimum temperatures of the inlet or hot spot temperature, then in the case of using MSA as the starting material, the amount of BSA increases, while the amount of GBL and BDO is reduced at the same time. Furthermore, deactivation of the catalyst by coating with succinic acid, fumaric acid and / or BSA and mechanical damage to the catalyst can be observed at such a temperature in the course of the hydrogenation.
  • the inlet temperature (initial hydrogenation temperature) is from 100 ° C. to 240 ° C., preferably 120 ° C. to 200 ° C., in particular 140 to 180 ° C. If the hydrogenation is carried out below the minimum temperatures of the inlet temperature, the yield of BDO decreases. The catalyst loses activity.
  • a pressure of 2 to 60 bar preferably a pressure of 2 to 20 bar and particularly preferably a pressure of 5 to 15 bar is selected.
  • this pressure range the formation of THF from the intermediate product GBL which is initially formed is largely suppressed in the hydrogenation range of MA.
  • a pressure of 20 to 250 bar preferably a pressure of 60 to 200 bar and particularly preferably a pressure of 80 to 160 bar is selected.
  • the conversion of GBL to BDO increases at the temperatures applied in the second hydrogenation stage.
  • a lower GBL recirculation rate can thus be selected by higher pressure.
  • the drain in the second hydrogenation stage is above the pressure in the first hydrogenation stage.
  • the catalyst loading of the first hydrogenation stage is preferably in the range from 0.02 to 1, in particular 0.05 to 0.5, kg of starting material / 1 catalyst • hour. If the catalyst load of the first stage in the case of MSA is increased beyond the range mentioned, an increase in the proportion of BSA and succinic acid in the hydrogenation discharge can generally be observed.
  • the catalyst loading of the second hydrogenation stage is preferably in the range from 0.02 to 1.5, in particular 0.1 to 1 kg of educt / 1 catalyst • hour. If the catalyst load is increased beyond the range mentioned, incomplete sales of GBL can be expected. This may be compensated for by an increased recycle rate, but this is of course not preferred.
  • the hydrogen / educt molar ratio is a parameter which has an influence on the product distribution and also on the economy of the process according to the invention.
  • a low hydrogen / starting material ratio is desirable from an economic point of view.
  • the lower limit is 5, but generally higher hydrogen / reactant molar ratios of 20 to 600 are used.
  • the use of the catalysts used in accordance with the invention and compliance with the temperature values applied in accordance with the invention permits the use of favorable, low hydrogen / starting material ratios in the hydrogenation of the first stage, which are preferably from 20 to 200, preferably from 40 to 150.
  • the cheapest range is from 50 to 100.
  • part, advantageously the main amount, of the hydrogen is usually circulated both in the first and in the second hydrogenation stage.
  • the cycle gas compressors known to the person skilled in the art are generally used.
  • the amount of hydrogen chemically consumed by the hydrogenation is supplemented.
  • part of the circulating gas is removed in order to remove inert compounds, for example n-butane.
  • the circulated hydrogen can also be used, if necessary after preheating, to evaporate the educt stream.
  • the cooling temperature is preferably 0 to 60 ° C, preferably 20 to 45 ° C.
  • reactor types all types of apparatus suitable for heterogeneously catalyzed reactions with a gaseous educt and product stream are considered as reactor types.
  • Tube-bundle reactors are particularly preferably used for the first hydrogenation stage.
  • the second hydrogenation is carried out as a suspension or fixed bed hydrogenation, preferably as a fixed bed hydrogenation.
  • All apparatuses suitable for heterogeneously catalyzed reactions in the liquid phase are suitable as reactor types for the second hydrogenation stage.
  • Tube reactors, loop reactors, stirred tanks, bubble columns, shaft reactors and reactors with internal heat removal, for example tube bundle reactors, are preferred. It is also possible to use a fluid bed. Loop or shaft reactors are particularly preferably used for the second hydrogenation stage.
  • reactors can be used in parallel or in series, both in the first and in the second hydrogenation stage.
  • an intermediate feed can take place between the catalyst beds. Intercooling between or in the catalyst beds is also possible.
  • the catalyst can be diluted with inert material.
  • An important point of the present invention is the choice of the catalysts for both stages, which have copper oxide as the catalytically active main component. This is attached to a carrier that may only have a small number of acidic centers. If a catalyst with an excessive number of acidic centers is used, BDO is dehydrated and THF is formed.
  • a suitable carrier material which has a sufficiently small number of acidic centers is, for example, a material selected from the group ZnO, Al 2 O 3 , SiO 2 , TiO 2 , ZrO 2 , CeO 2 , MgO, CaO, SrO, coal of various origins , BaO and Mn 2 O 3 and mixtures thereof.
  • Preferred carrier materials are ZnO / Al 2 O 3 mixtures, the delta, theta, alpha and eta modifications of Al 2 O 3 and mixtures which contain at least one component from the group Si0 2 , TiO, ZrO 2 on the one hand and from the group ZnO, MgO, CaO, SrO and BaO on the other hand.
  • the amount of copper oxide is ⁇ 95% by weight, preferably 5 to 95% by weight, in particular 15 to 80% by weight; the carrier is used in amounts of> 5% by weight, preferably 5 to 95% by weight, in particular 20 to 85% by weight. Because of the toxicity of chromium-containing catalysts, chromium-free catalysts are preferably used. Of course, corresponding chromium-containing catalysts known to the person skilled in the art are also technically suitable for use in the process according to the invention, but this does not result in the desired advantages, which are in particular environmental and work-related in nature.
  • the same catalyst can be used in both hydrogenation stages, but the use of different catalysts is preferred.
  • the catalysts used according to the invention can contain one or more further metals or a compound thereof, preferably an oxide, from groups 1 to 14 (IA to VIDA and IB to IVB of the old IUPAC nouns) of the Periodic Table of the Elements.
  • Pd is preferably used in amounts of ⁇ 1% by weight, preferably ⁇ 0.5% by weight, in particular ⁇ 0.2% by weight.
  • the addition of another metal or metal oxide is not preferred.
  • the catalysts used can also contain an auxiliary in an amount of 0 to 10% by weight.
  • auxiliary are understood to be organic and inorganic substances which contribute to improved processing during catalyst production and / or to an increase in the mechanical strength of the shaped catalyst bodies. Such aids are known to the person skilled in the art; Examples include cement, graphite, stearic acid, silica gel and copper powder.
  • the catalysts can be prepared by methods known to those skilled in the art. Preference is given to processes in which the copper oxide is finely divided and intimately mixed with the other constituents; impregnation and precipitation reactions are particularly preferred.
  • These starting materials can be processed to shaped articles by known methods, for example extrusion, tableting or by agglomeration processes, if appropriate with the addition of auxiliaries.
  • catalysts according to the invention can also be produced, for example, by applying the active component to a support, for example by coating or vapor deposition. Furthermore, catalysts according to the invention can be used Deforming a heterogeneous mixture of active component or precursor compound thereof with a carrier component or precursor compound thereof can be obtained.
  • the catalyst is used in reduced, activated form.
  • the activation is carried out using reducing gases, preferably hydrogen or hydrogen / inert gas mixtures, either before or after installation in the reactors in which the process according to the invention is carried out. If the catalyst was installed in the reactor in oxidic form, the activation can be carried out both before starting up the plant with the hydrogenation according to the invention and during start-up, that is to say in situ.
  • the separate activation before starting up the system is generally carried out using reducing gases, preferably hydrogen or hydrogen / inert gas mixtures, at elevated temperatures, preferably between 100 and 350 ° C. In the so-called in-situ activation, the activation takes place when the system is started up by contact with hydrogen at an elevated temperature.
  • the catalysts are preferably used as moldings. Examples include strands, rib strands, other extras, tablets, rings, balls and grit.
  • the BET surface area of the copper catalysts should generally be 10 to 1500 m 2 / g in the oxidic state, preferably 10 to 300 m 2 / g, more preferably 15 to 175 m 2 / g, in particular 20 to 150 m 2 / g.
  • the copper surface (N 2 O decomposition) of the reduced catalyst should be> 0.2 m Ig, preferably> 1 m / g, in particular> 2 m / g in the installed state.
  • catalysts which have a defined porosity.
  • these catalysts have a pore volume of> 0.01 ml / g for pore diameters> 50 nm, preferably> 0.025 ml / g for pore diameters> 100 nm and in particular> 0.05 ml / g for pore diameters> 200 nm
  • the porosities mentioned were determined by mercury intrusion according to DEN 66133.
  • the data in the pore diameter range from 4 nm to 300 ⁇ m were evaluated.
  • the catalysts used according to the invention generally have a sufficient service life. In the event that the activity and / or selectivity of the catalyst should nevertheless decrease in the course of its operating time, it can be regenerated by measures known to the person skilled in the art.
  • This preferably includes reductive treatment of the catalyst in a hydrogen stream at elevated temperature. If necessary, the reductive treatment can be preceded by an oxidative treatment. In this case, a molecular mixture containing gas, for example air, flows through the catalyst bed at elevated temperature. It is also possible to wash the catalysts with a suitable solvent, for example ethanol, THF, BDO or GBL, and then to dry them in a gas stream.
  • a suitable solvent for example ethanol, THF, BDO or GBL
  • a mixed basic metal carbonate is precipitated from a metal salt solution containing copper nitrate and zinc nitrate with soda at 50 ° C. and a pH around 6.2.
  • the metal salt solution used contained the metals corresponding to a catalyst composition of 70% CuO and 30% ZnO.
  • the precipitate is filtered, washed, dried, calcined at 300 ° C. and compressed with 3% by weight of graphite to give tablets of 3 mm in height and diameter.
  • the catalyst Before the start of the reaction, the catalyst is subjected to a hydrogen treatment in the hydrogenation apparatus.
  • the reactor is heated to 180 ° C. and the catalyst is activated for the time specified in Table 1 with the mixture of hydrogen and nitrogen given at atmospheric pressure.
  • the pressure equipment used for the hydrogenation consists of an evaporator, a reactor, a cooler with quench feed, a hydrogen supply, an exhaust pipe and a circulating gas blower.
  • the drain in the apparatus is kept constant.
  • the melted MSA is pumped from above onto the preheated (245 ° C) evaporator and evaporated.
  • a mixture of fresh hydrogen and circulating gas also reaches the evaporator from above.
  • Hydrogen and MSA thus enter the temperature-controlled reactor from below.
  • the reactor contents consist of a mixture of glass rings and catalyst.
  • the condensed liquid reaction discharge, the exhaust gas and the cycle gas are analyzed quantitatively by gas chromatography.
  • a solid with the composition 64% ZnO and 36% Al 2 O 3 (based on 100% oxide) is precipitated from an aqueous solution of zinc nitrate and aluminum nitrate with sodium carbonate solution at 50 ° C. and a pH of 6.8, filtered and washed. The filter cake was dried and calcined at 425 ° C for one hour.
  • the carrier described above is added to a nitric acid solution of copper nitrate and zinc nitrate (metal ratio corresponding to 16.6% by weight CuO and 83.4% ZnO) and mixed intensively at 70 ° C.
  • a solid is precipitated from this mixture with sodium carbonate solution at 70 ° C. and a pH of 7.4 and the suspension is left to stir for 2 hours at constant temperature and pH.
  • the solid is filtered off, washed, dried and calcined at 430 ° C. for one hour.
  • the catalyst powder obtained in this way was mixed with 1.5% by weight of graphite and 5% by weight of copper powder and pressed into tablets of 1.5 mm in diameter and 1.5 mm in height.
  • the tablets were finally calcined at 330 ° C. for 1 hour, had a lateral tensile strength of 50N and a chemical composition of 66% CuO / 24% ZnO / 5% Al 2 O 3 /5% Cu.
  • catalyst activation analogous to example lb
  • the catalyst described in example 2a is filled in a 300 ml autoclave and the reaction discharge from the MSA hydrogenation in example 1 is added. After a reaction time of 48 h at 180 ° C and 80 bar 89% BDO, 2% GBL, 6% THF could be detected in the discharge.

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  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
  • Catalysts (AREA)
  • Low-Molecular Organic Synthesis Reactions Using Catalysts (AREA)

Abstract

L'invention concerne un procédé de production de 1,4-butanediol éventuellement substitué par alkyle, par hydrogénation catalytique en deux étapes, dans la phase gazeuse d'acides dicarboxyliques C4 et/ou de leurs dérivés, ledit procédé comprenant les étapes suivantes: a) introduction d'un flux gazeux d'un acide dicarboxylique C4 ou d'un dérivé de ce dernier dans un premier réacteur et hydrogénation catalytique pour l'obtention d'un produit contenant principalement de la gamma-butyrolactone éventuellement substituée par alkyle ; b) conversion du flux de produit en phase liquide; c) introduction du flux de produit ainsi obtenu dans un deuxième réacteur et hydrogénation catalytique dans la phase liquide pour l'obtention du 1,4-butanediol; d) séparation des produits secondaires et éventuellement de l'éduit n'ayant pas réagi, présents dans le produit voulu; un catalyseur est utilisé lors des deux étapes d'hydrogénation, ledit catalyseur contenant ≤ 95 % en volume d'un support. Dans le deuxième réacteur, la pression est plus élevée que dans le premier, et le mélange de produit prélevé dans le premier réacteur est introduit dans le deuxième réacteur sans épuration supplémentaire.
EP03740211A 2002-06-11 2003-06-10 Procede pour produire du 1,4-butanediol par hydrogenation combinee en phase gazeuse et en phase liquide Withdrawn EP1515935A1 (fr)

Applications Claiming Priority (3)

Application Number Priority Date Filing Date Title
DE10225927 2002-06-11
DE10225927A DE10225927A1 (de) 2002-06-11 2002-06-11 Verfahren zur Herstellung von Butandiol durch kombinierte Gasphasen- und Flüssigphasensynthese
PCT/EP2003/006057 WO2003104176A1 (fr) 2002-06-11 2003-06-10 Procede pour produire du 1,4-butanediol par hydrogenation combinee en phase gazeuse et en phase liquide

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EP1515935A1 true EP1515935A1 (fr) 2005-03-23

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EP (1) EP1515935A1 (fr)
JP (1) JP2005534658A (fr)
KR (1) KR100898455B1 (fr)
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AU (1) AU2003274707A1 (fr)
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GB201318175D0 (en) 2013-10-14 2013-11-27 Johnson Matthey Davy Technologies Ltd Process
CN106311255B (zh) * 2015-07-03 2019-12-24 中国石油化工股份有限公司 一种制备丁二醇用催化剂及其制备方法和应用
CN107537516B (zh) * 2016-06-28 2020-05-12 中国石油化工股份有限公司 用于制备仲丁醇的催化剂及其制备方法
CN112500418A (zh) * 2020-12-31 2021-03-16 山东中科新材料研究院有限公司 一种采用回路反应器合成1，8-二氮杂二环十一碳-7-烯的方法

Family Cites Families (60)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
AT257556B (de) 1962-11-16 1967-10-10 Celanese Corp Verfahren zur Herstellung von Alkoholen
FR1558895A (fr) 1967-04-11 1969-02-28
GB1230276A (fr) 1968-12-09 1971-04-28
GB1344557A (en) 1972-06-23 1974-01-23 Mitsubishi Petrochemical Co Process for preparing 1,4-butanediol
JPS5024285B2 (fr) 1972-06-26 1975-08-14
DE2231986C2 (de) 1972-06-29 1975-01-09 Mitsubishi Petrochemical Co. Ltd., Tokio Verfahren zur herstellung von 1,4 butandiol
JPS565210B2 (fr) 1972-12-27 1981-02-04
US4001282A (en) 1973-02-12 1977-01-04 General Electric Company Process for producing gamma-butyrolactone
PL87238B1 (fr) 1973-07-28 1976-06-30
CA1020673A (fr) 1973-08-30 1977-11-08 Kazuyoshi Tsukamoto Bloc d'accord automatique
DE2455617C3 (de) 1974-11-23 1982-03-18 Basf Ag, 6700 Ludwigshafen Verfahren zur Herstellung von Butandiol und/oder Tetrahydrofuran über die Zwischenstufe des γ-Butyrolactons
DE2501499A1 (de) 1975-01-16 1976-07-22 Hoechst Ag Verfahren zur herstellung von butandiol-(1.4)
DE2845905C3 (de) 1978-10-21 1983-05-26 Chemische Werke Hüls AG, 4370 Marl Verfahren zur kontinuierlichen Herstellung von Butandiol-1,4
US4301077A (en) 1980-12-22 1981-11-17 Standard Oil Company Process for the manufacture of 1-4-butanediol and tetrahydrofuran
IT1190783B (it) 1981-04-29 1988-02-24 Davy Mckee Oil & Chem Processo per l'idrogenolisi di esteri di acidi carbossilici
US4652685A (en) 1985-11-15 1987-03-24 General Electric Company Hydrogenation of lactones to glycols
DE3726510A1 (de) 1987-08-08 1989-02-16 Basf Ag Verfahren zur herstellung von tetrahydrofuran oder von gemischen, die neben tetrahydrofuran (gamma)-butyrolacton und/oder 1,4-butandiol enthalten
EP0304696B1 (fr) 1987-08-08 1992-04-29 BASF Aktiengesellschaft Procédé pour la préparation de 1,4-butanediol et/ou tétrahydrofuranne
JPH07116082B2 (ja) 1987-11-04 1995-12-13 三菱化学株式会社 １，４−ブチレングリコールの製造法
US4797382A (en) 1987-11-27 1989-01-10 Gaf Corporation Hydrogenation catalyst and process for preparing the catalyst
US4885411A (en) 1987-11-27 1989-12-05 Gaf Corporation Hydrogenation process for converting lactones to diols
US4965378A (en) 1987-12-23 1990-10-23 The Standard Oil Company Vapor-phase hydrogenation of maleic anhydride to tetrahydrofuran and gamma-butyrolactone
US5072009A (en) 1987-12-23 1991-12-10 The Standard Oil Company Vapor-phase hydrogenation of maleic anhydride to tetrahydrofuran and gamma-butyrolactone
CA1327812C (fr) 1987-12-23 1994-03-15 Thomas G. Attig Hydrogenation en phase vapeur de l'anhydride maleique en tetrahydrofurane et en gamma-butyrolactone
JP2772524B2 (ja) 1988-07-15 1998-07-02 東燃株式会社 １，４‐ブタンジオールの製造方法
JP2666151B2 (ja) 1988-12-14 1997-10-22 東燃株式会社 １，４−ブタンジオールの製造法
JP2596604B2 (ja) 1988-12-14 1997-04-02 東燃株式会社 １，４−ブタンジオールおよびテトラヒドロフランの製造方法
DE3903363A1 (de) 1989-02-04 1990-08-23 Basf Ag Verfahren zur herstellung von alkanolen
JP2639462B2 (ja) 1989-03-08 1997-08-13 東燃株式会社 １，４―ブタンジオールおよびテトラヒドロフランの製造法
JP2639464B2 (ja) 1989-03-08 1997-08-13 東燃株式会社 １，４―ブタンジオールおよびテトラヒドロフランの製法
JP2639463B2 (ja) 1989-03-08 1997-08-13 東燃株式会社 １，４―ブタンジオールおよびテトラヒドロフランの製造方法
JP2670698B2 (ja) 1989-03-08 1997-10-29 東燃株式会社 １，４―ブタンジオールの製法
US5055599A (en) 1989-06-23 1991-10-08 The Standard Oil Company Process for the hydrogenation of maleic anhydride to tetrahydrofuran and gamma-butyrolactone
JP2595358B2 (ja) 1989-12-07 1997-04-02 東燃株式会社 １，４―ブタンジオールおよびテトラヒドロフランの製造方法
US5347021A (en) 1990-04-16 1994-09-13 Isp Investments Inc. Process of vapor phase catalytic hydrogenation of maleic anhydride to gamma-butyrolactone in high conversion and high selectivity using an activated catalyst
CN1015901B (zh) 1990-07-10 1992-03-18 中国石油化工总公司 气控内循环式催化剂冷却器
US5030773A (en) 1990-07-25 1991-07-09 General Electric Company Process for the production of butanediol
US5149836A (en) 1991-07-25 1992-09-22 Isp Investments Inc. Process for the production of gamma butyrolactone THF in predetermined amounts
US5122495A (en) 1991-10-02 1992-06-16 Isp Investments Inc. Activated catalyst for the vapor phase hydrogenation of maleic anhydride to gamma-butyrolactone in high conversion and high selectivity
US5196602A (en) * 1991-12-30 1993-03-23 The Standard Oil Company Two-stage maleic anhydride hydrogenation process for 1,4-butanediol synthesis
JPH0634567A (ja) 1992-07-17 1994-02-08 Murata Mfg Co Ltd ワーク欠損検出装置
DE4326692A1 (de) 1993-08-10 1995-02-16 Akzo Nobel Nv Verfahren zur Herstellung von gamma-Butyrolacton
GB9324753D0 (en) 1993-12-02 1994-01-19 Davy Mckee London Process
GB9324782D0 (en) 1993-12-02 1994-01-19 Davy Mckee London Process
IT1273320B (it) 1994-02-22 1997-07-08 Alusuisse Italia Spa Procedimento per la produzione di gamma-butirrolattone
CN1046216C (zh) 1994-06-04 1999-11-10 中国石油化工总公司 制备1,4-丁二醇和/或γ-丁内酯的催化剂
CN1049207C (zh) 1994-08-10 2000-02-09 中国石油化工总公司 １，４－丁二醇的制备方法
CN1046434C (zh) 1995-06-14 1999-11-17 中国石油化工总公司 气相氢化制1,4－丁二醇的催化剂(二)
CN1052663C (zh) 1995-06-14 2000-05-24 中国石油化工总公司 气相氢化制1,4－丁二醇的催化剂
TW341568B (en) 1995-12-27 1998-10-01 Akzo Nobel Nv Process for manufacturing Gamma-butyrolactone and its use
ZA973971B (en) 1996-05-15 1998-03-23 Kvaerner Process Tech Ltd A process for the production of at least one C4 compound selected from butane-1,4-diol, gamma-butyrolactone and tetrahydrofuran.
CN1059194C (zh) 1996-11-15 2000-12-06 中国石油化工总公司 一种气相加氢制1,4-丁二醇的方法
CN1072524C (zh) 1996-11-15 2001-10-10 中国石油化工总公司 一种气相加氢制1,4－丁二醇的催化剂
BE1011698A6 (fr) 1998-01-08 1999-12-07 Pantochim Sa Procede de production de tetrahydrofuranne, de gamma-butyrolactone et de butanediol.
BE1011699A6 (fr) 1998-01-08 1999-12-07 Pantochim Sa Procede de production de tetrahydrofuranne et de gamma-butyrolactone.
BE1012342A6 (fr) 1998-01-08 2000-10-03 Pantochim Sa Procede de production de butanediol par hydrogenation en phase liquide.
IT1298096B1 (it) 1998-01-09 1999-12-20 Lonza Spa Procedimento per la produzione di gamma-butirrolattone
IT1298535B1 (it) 1998-02-02 2000-01-12 Lonza Spa Procedimento per la produzione di gamma-butirrolattone
US6008384A (en) * 1998-03-03 1999-12-28 E. I. Du Pont De Nemours And Company Method and Ru,Re,Sn/carbon catalyst for hydrogenation in aqueous solution
BE1011870A6 (fr) 1998-04-09 2000-02-01 Pantochim Sa Procede de production de butanediol.

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
See references of WO03104176A1 *

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DE10225927A1 (de) 2003-12-24
US7169958B2 (en) 2007-01-30
CN1294110C (zh) 2007-01-10
US20050267318A1 (en) 2005-12-01
WO2003104176A1 (fr) 2003-12-18
JP2005534658A (ja) 2005-11-17
KR20050007606A (ko) 2005-01-19
KR100898455B1 (ko) 2009-05-21
CN1659124A (zh) 2005-08-24
AU2003274707A1 (en) 2003-12-22

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