200408622 玖、發明說明： 【發明所屬之技術領域】 本發明關於一種製造芳香族碳酸酯之方法。確切地，本 發明關於一種經由碳酸二芳酯與芳族羥基化合物之轉酯化 反應而有效率地及連續地製造碳酸烷基芳酯及/或碳酸二 芳酯之方法。 【先前技術】 迄今已熟知經由碳酸二芳酯與芳族羥基化合物之轉酯 化反應製造碳酸烷基芳酯以及自碳酸烷基芳酯製造碳酸二 芳酯。針對此等反應可以下式（1 )至（3 )表示： nCOO-RHAfOHUCOO - Ar + R20H (1) R丨-OCOO-Ar + ArOH — Ar-OCOO-Ar + R 丨 OH (2) 2R1-0C00-Ar->Ar-0C00-Ar+R1-0C00-R1 (3) (其中R 1及R2每一代表脂族烴基團或脂環族烴基團，且其 可相同或不同；且Ar代表芳族烴基團。） 此類型轉酯化反應係為平衡反應，且其具有朝強親核取 代基取代弱親核取代基方向進行之趨向。假如於起始碳酸 二烷酯為具較低脂族烴基團脂的碳酸二烷酯且其他起始芳 族經基化合物為盼之場合中，兩者之反應強烈地傾向起始 系統，因為式（1 )與（2 )之反應與理論相反，且反應速率通 常是低的。當使用通常的轉酯化反應觸媒（例如鹼金屬氫氧 化物）時，則具下式（4 )之反應（但不是式（1 )之反應）將提高 為支配地位，因此，反應產率將明顯地降低。 R1-0C00-R2+Ar0H->R1-0-Ar+R20H+C02 (4) 5 312/發明說明書(補件)/92-10/92120780 200408622 為了有效率地促進式（1)、（2)及（3)之反應，先前已尋 找具高活性觸媒，且已提出許多觸媒。舉例來說，使用含 任一種有機錫化合物（日本專利特許公開申請案第 4 8 7 3 3 / 1 9 7 9號）或有機鈦化合物（日本專利特許公開申請 案第1 8 3 7 4 5 / 1 9 8 2號）之絡合物觸媒可促進上述反應。 為了更有效率地製造芳族碳酸酯，必須快速地自反應系 統移走產物，俾平衡可儘可能地朝向產物系統。為了有效 率地移走副產物，舉例來說，經揭示自反應系統移走脂族 醇者為一種經由共沸蒸餾將其移走之方法（日本專利特許 公開申請案第4 8 7 3 2 / 1 9 7 9或2 9 1 5 4 5 / 1 9 8 6號）、一種使用 分子篩將其吸附及移走之方法（日本專利特許公開申請案 第1 8 5 5 3 6 / 1 9 8 3號）及一種利用滲透蒸發或蒸汽蒸發之方 法（曰本專利特許公開申請案第1 2 5 0 2 1 / 1 9 9 3號）。然而， 由於其缺點為實驗規模不易按比例放大且此等方法複雜， 故全部不適用於工業方法。 於式（1)之平衡反應中，起始材料之轉化率不能超過連 續攪拌槽反應器中之平衡組成物。為了解決此問題，使用 多階段反應器串聯連接之反應法可能是有效的，其中係以 逐步增加轉化率自每一階段取出產物。基於相同的理由， 已知可連續地取出產物之反應性蒸餾作用將亦是有效的。 舉例來說，於日本專利公告第9 1 2 3 6 / 1 9 9 5號中，液態芳族 羥基化合物經供給至連續多階段蒸餾塔中（經由其頂部）， 同時氣態碳酸二烷酯係於其中通過底部，以致於兩者於塔 中彼此朝逆流接觸，且含有芳族碳酸酯之高沸點成分係經 6 312/發明說明書(補件)/92-10/92120780 200408622 由蒸餾塔底部取出，同時將含有副產物醇及碳酸二烷酯之 低沸點成分經由其頂部取出。 理論上，此方法之反應可以相當簡單的程序進行，但反 應速率是緩慢的。此外，由於反應為液相反應，故充足的 反應時間不易確保其在此一連續多階段蒸餾塔中。為了確 保其必須的反應時間，已揭示一種附加一個額外反應器之 方法（日本專利特許公開申請案第2 2 4 5 4 7 / 1 9 9 2或 2 3 0 2 4 2 / 1 9 9 2號），但供其使用之修飾的設備將不可避免地 複雜及昂貴。 已揭示一種在至少兩個串聯連接的攪拌槽中以連續蒸 氣-液體逆流流體進行反應之方法（日本專利特許公開申請 案第2 3 4 7 0 7 / 1 9 9 4號）及一種於發泡塔反應器中或於至少 二個以串聯方式連接之發泡塔中進行反應之方法（日本專 利特許公開申請案第2 9 8 7 0 0 / 1 9 9 4號），因為此等方法可獲 致與反應性蒸餾相同的結果。此等方法是有利的，因為反 應時間可自由地改變，但不利的是，倘若反應階段數目少 的時候無法得到充足的轉化率，但倘若反應階段數目大的 時候將提高設備成本。 一般而言，多階段反應係於反應性蒸餾類型反應器裝置 中進行。具體地，於較低階段或較後階段中產生之蒸氣連 續地朝逆流接觸，且與來自較高階段或較前階段之反應媒 介物反應，蒸氣-液體反應經進行，且有利的是供應至較低 階段或較後階段之能量可有利地轉移至較高階段或較前階 段。就另一方面而論，低沸點產物（於此情形中為脂族醇） 7 312/發明說明書(補件)/92-10/92120780 200408622 在較高階段或較前階段中冷凝更多，同時高沸點產物（於此 情形中為芳族碳酸酯）在較低階段或較後階段中冷凝更 多，此外，起始化合物之比例連續地改變。因此，反應不 易控制，且反應器規劃及操作通常面臨困難。 就移走多階段反應之每一階段中製得的脂族醇而言，並 非總是需要反應性蒸餾類型裝置。然而，倘若在高沸點產 物移走後蒸氣所具有的能量未有效率地自反應系統回收， 則能量效率將極端地變差。 就另一方面而論，當攪拌槽串聯連接時，則設備成本將 如上述日本專利特許公開申請案第2 3 4 7 0 7 / 1 9 9 4號中之方 法般地提高。 為了降低多階段反應中之設備成本，建議在反應器中進 行之反應方法為其中藉隔牆分開液相部分，蒸氣相形成連 續相，且經由反應器頂部連續地取出含氣相副產物（脂族醇） 之高沸點部分（曰本專利特許公開申請案第1 8 8 5 5 8 / 1 9 9 6 號）。於此方法中，可改變個別分隔反應區中之液相的溫度 及組成，且反應可以多階段反應進行。然而，由於其中每 一反應區形成連續相，故此方法仍有缺陷，因為根據其中 之蒸氣組成，將無法達成適當的分離及回收處理及能量回 收。 如上述，經由傳統轉酯化反應製造碳酸烷基芳酯及/或 碳酸二芳酯之方法具有一些缺點，因為其不易達到具高選 擇率及高效率之良好反應，且其需要複雜的步驟及昂貴的 設備。 8 312/發明說明書(補件)/92-10/92120780 200408622 本發明之目的係提供一種於觸媒存在下自碳酸二烷酯 與芳族羥基化合物製造碳酸烷基芳酯及/或碳酸二芳酯之 方法，此方法沒有上述缺點，且使具選擇性、有效率及連 續製造所欲產物成為可能。 【發明内容】 本案發明人已鍥而不捨地研究解決上述問題，因而有如 下發現：轉酯化反應係為液相平衡反應，因此，藉有效率 地移走低沸點產物（脂族醇）可妥善地促進反應；因為當反 應達到更平衡時，反應促進延遲，多階段反應進行（針對每 一階段中改變之條件）；且於一階段中形成之蒸氣不會直接 地與反應液體接觸；但經由間接熱交換，其用於加熱此階 段中或前一階段中之反應液體，因而改良能量效率及反應 延遲（由於可避免低沸點副產物之緣故），且因此，此方式 之受控方法較反應性蒸餾更有利。於此等發現為基礎，本 案發明人已完成本發明。 特別地，本發明關於一種製造芳族碳酸酯之方法，製造 碳酸烷基芳酯及/或碳酸二芳酯之方法係藉著使碳酸二烷 酯與芳族羥基化合物於觸媒存在下反應，轉酯化反應係於 具至少二個串聯連接的獨立反應區結構之反應裝置中進 行，依此一受控方式，將液相依序引導進入第一反應區至 最終反應區，且在至少一反應區中產生的蒸氣之冷凝熱係 轉移至供給至該反應區或進入前一階段反應區之液體，以 便加熱該液體，並且將蒸氣保持離開以免直接接觸所欲的 反應液體。 9 312/發明說明書(補件)/92-10/9212〇780 200408622 【實施方式】 以下將詳細地說明進行本發明之方式。 於本發明中之起始材料碳酸二烷酯係以下式（5 )表示： Ri-OCOO-R2 (5) (其中R1及R2每一代表具1至10個碳原子之烷基基團， 且R1及R2可相同或相異。） 具體地，其包含碳酸二甲酯、碳酸二乙酯、碳酸二丙酯、 碳酸二丁酯、碳酸甲基乙酯。當然，特佳者為碳酸二曱酯 及碳酸二乙酯。 於本發明中之另一起始材料芳族羥基化合物係以下式 (6 )表示：200408622 发明, Description of the invention: [Technical field to which the invention belongs] The present invention relates to a method for manufacturing an aromatic carbonate. Specifically, the present invention relates to a method for efficiently and continuously producing an alkylaryl carbonate and / or a diaryl carbonate via a transesterification reaction of a diaryl carbonate and an aromatic hydroxy compound. [Prior art] Hitherto, it has been well known to produce an alkylaryl carbonate through a transesterification reaction of a diaryl carbonate and an aromatic hydroxy compound, and to produce a diaryl carbonate from an alkylaryl carbonate. For these reactions, the following formulas (1) to (3) can be expressed: nCOO-RHAfOHUCOO-Ar + R20H (1) R 丨 -OCOO-Ar + ArOH — Ar-OCOO-Ar + R 丨 OH (2) 2R1-0C00 -Ar- > Ar-0C00-Ar + R1-0C00-R1 (3) (where R 1 and R 2 each represent an aliphatic hydrocarbon group or an alicyclic hydrocarbon group, and they may be the same or different; and Ar represents an aromatic group Group hydrocarbon groups.) This type of transesterification reaction is an equilibrium reaction, and it has a tendency to replace strong nucleophilic substituents with strong nucleophilic substituents. If the starting dialkyl carbonate is a dialkyl carbonate with a lower aliphatic hydrocarbon group lipid and other starting aromatic compounds are desired, the reaction between the two strongly favors the starting system because the formula The reaction of (1) and (2) is contrary to theory, and the reaction rate is usually low. When a common transesterification reaction catalyst (such as an alkali metal hydroxide) is used, the reaction of the following formula (4) (but not the reaction of the formula (1)) will be dominated. Therefore, the reaction yield Will be significantly reduced. R1-0C00-R2 + Ar0H- > R1-0-Ar + R20H + C02 (4) 5 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 In order to efficiently promote equations (1), (2 ) And (3), many catalysts have been previously searched for, and many catalysts have been proposed. For example, use any organic tin compound (Japanese Patent Laid-open Application No. 4 8 7 3 3/1 9 7 9) or an organic titanium compound (Japanese Patent Laid-Open Application No. 1 8 3 7 4 5 / The complex catalyst of No. 9 8 2) can promote the above reaction. In order to produce aromatic carbonates more efficiently, the products must be removed from the reaction system quickly, and the tritium equilibrium can be directed towards the product system as much as possible. In order to efficiently remove by-products, for example, it has been disclosed that removing aliphatic alcohols from the reaction system is a method of removing them by azeotropic distillation (Japanese Patent Laid-Open Application No. 4 8 7 3 2 / 1 9 7 9 or 2 9 1 5 4 5/1 9 8 6), a method for adsorbing and removing molecular sieves (Japanese Patent Laid-open Application No. 1 8 5 5 3 6/1 9 8 3 ) And a method using pervaporation or steam evaporation (this patent application publication No. 1 2 50 2 1/1 193). However, due to the disadvantages that the scale of experiments is not easy to scale up and that these methods are complicated, they are not suitable for industrial methods. In the equilibrium reaction of formula (1), the conversion rate of the starting material cannot exceed the equilibrium composition in the continuous stirred tank reactor. To solve this problem, it may be effective to use a multi-stage reactor connected in series, in which the product is taken out from each stage with a stepwise increase in conversion. For the same reason, it is known that reactive distillation which can continuously take out the product will also be effective. For example, in Japanese Patent Publication No. 9 1 2 3 6/19 95, the liquid aromatic hydroxy compound is supplied to a continuous multi-stage distillation column (through its top), and the gaseous dialkyl carbonate is at the same time Wherein through the bottom, so that the two are in countercurrent contact with each other in the column, and the high boiling point component containing aromatic carbonate is taken out from the bottom of the distillation column via 6 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622, At the same time, low-boiling components containing by-product alcohol and dialkyl carbonate were taken out through the top. Theoretically, the reaction of this method can be performed with a fairly simple procedure, but the reaction rate is slow. In addition, since the reaction is a liquid phase reaction, it is not easy to ensure a sufficient reaction time in this continuous multi-stage distillation column. In order to ensure the necessary reaction time, a method for attaching an additional reactor (Japanese Patent Laid-Open Application No. 2 2 4 5 4 7/1 9 9 2 or 2 3 0 2 4 2/1 9 9 2 ), But the modified equipment for its use will inevitably be complicated and expensive. A method for reacting with a continuous vapor-liquid countercurrent fluid in at least two stirred tanks connected in series has been disclosed (Japanese Patent Laid-Open Application No. 2 3 4 7 0 7/1 9 9 4) and a method for foaming A method for carrying out a reaction in a column reactor or in at least two foaming towers connected in series (Japanese Patent Laid-Open Application No. 2 987 0 0/19 9 4), because these methods can be obtained Same result as reactive distillation. These methods are advantageous because the reaction time can be freely changed, but disadvantageously, if the number of reaction stages is small, a sufficient conversion rate cannot be obtained, but if the number of reaction stages is large, the equipment cost will be increased. Generally, a multi-stage reaction is performed in a reactive distillation type reactor apparatus. Specifically, the vapor generated in the lower or later stage is continuously contacted countercurrently and reacts with the reaction medium from the higher or earlier stage. The vapor-liquid reaction proceeds and is advantageously supplied to The energy of the lower or later stages can be advantageously transferred to the higher or earlier stages. On the other hand, low-boiling products (aliphatic alcohols in this case) 7 312 / Explanation of the Invention (Supplement) / 92-10 / 92120780 200408622 Condensate more in the higher stage or in the previous stage, while The high-boiling product (aromatic carbonate in this case) condenses more in the lower or later stages, and in addition, the proportion of the starting compound changes continuously. Therefore, the reaction is not easy to control, and reactor planning and operation are often difficult. To remove the aliphatic alcohol produced in each stage of a multi-stage reaction, a reactive distillation type device is not always required. However, if the energy of the vapor is not efficiently recovered from the reaction system after the high-boiling products are removed, the energy efficiency will be extremely deteriorated. On the other hand, when the mixing tanks are connected in series, the equipment cost will increase as in the method described in the aforementioned Japanese Patent Laid-open Application No. 2 3 4 0 7/19 4. In order to reduce the equipment cost in the multi-stage reaction, the reaction method proposed in the reactor is to separate the liquid phase part by a partition wall, the vapor phase forms a continuous phase, and the gas phase by-products (lipids) are continuously taken out through the top of the reactor. Group alcohols) (the Japanese Patent Laid-open Application No. 1 8 8 5 5/19 96). In this method, the temperature and composition of the liquid phase in the individual partitioned reaction zones can be changed, and the reaction can be carried out in multiple stages. However, since a continuous phase is formed in each of the reaction zones, this method still has disadvantages, because according to the vapor composition therein, proper separation and recovery processing and energy recovery cannot be achieved. As described above, the method for producing alkylaryl carbonate and / or diaryl carbonate through the traditional transesterification reaction has some disadvantages, because it is not easy to achieve a good reaction with high selectivity and high efficiency, and it requires complicated steps and Expensive equipment. 8 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 The object of the present invention is to provide an alkylaryl carbonate and / or diaryl carbonate produced from a dialkyl carbonate and an aromatic hydroxy compound in the presence of a catalyst. Ester method, this method does not have the above disadvantages, and makes it possible to selectively, efficiently and continuously produce the desired product. [Summary of the Invention] The inventors of this case have persistently researched and solved the above-mentioned problems, and thus have the following findings: The transesterification reaction is a liquid phase equilibrium reaction. Therefore, by efficiently removing the low boiling point product (aliphatic alcohol), it can be properly promoted Reaction; because when the reaction reaches a more equilibrium, the reaction promotion is delayed, the multi-stage reaction proceeds (for the conditions changed in each stage); and the vapor formed in one stage does not directly contact the reaction liquid; but via indirect heat Exchange, which is used to heat the reaction liquid in this stage or in the previous stage, thereby improving energy efficiency and reaction delay (because low-boiling by-products can be avoided), and therefore, the controlled method in this way is more reactive than distillation More favorable. Based on these findings, the present inventors have completed the present invention. In particular, the present invention relates to a method for producing an aromatic carbonate, and a method for producing an alkylaryl carbonate and / or a diaryl carbonate by reacting a dialkyl carbonate with an aromatic hydroxy compound in the presence of a catalyst, The transesterification reaction is performed in a reaction device having at least two independent reaction zone structures connected in series. In this controlled manner, the liquid phase is sequentially guided into the first reaction zone to the final reaction zone, and at least one The heat of condensation of the vapor generated in the reaction zone is transferred to the liquid supplied to the reaction zone or into the reaction zone of the previous stage in order to heat the liquid and keep the vapor away from direct contact with the desired reaction liquid. 9 312 / Invention Specification (Supplement) / 92-10 / 9212〇780 200408622 [Embodiment] Hereinafter, a mode for carrying out the present invention will be described in detail. The starting material dialkyl carbonate in the present invention is represented by the following formula (5): Ri-OCOO-R2 (5) (where R1 and R2 each represent an alkyl group having 1 to 10 carbon atoms, and R1 and R2 may be the same or different.) Specifically, they include dimethyl carbonate, diethyl carbonate, dipropyl carbonate, dibutyl carbonate, and methyl ethyl carbonate. Of course, particularly preferred are dimethyl carbonate and diethyl carbonate. The aromatic hydroxy compound as another starting material in the present invention is represented by the following formula (6):
ArOH (6) (其中A r代表具6至2 0個碳原子之芳族基團。） 具體地，其包含酚、鄰-甲酚、間'曱酚或對-甲酚、鄰-乙基酚、間-乙基酚或對-乙基酚、鄰-丙基酚、間-丙基酚 或對-丙基酚、鄰-曱氧基酚、間-甲氧基酚或對-甲氧基酚、 2, 6-二甲基酚、2,4 -二甲基酚、3, 4-二甲基酚、鄰-氯酚、 間-氯酚或對-氯酚、1 -萘酚、2 -萘酚。其中特佳者為酚。 本發明方法中之一產物碳酸烷基芳酯係以下式（7 )表 示： R3-0C00-Ar (7) (其中Ar具有與式（6)中相同的意義；且R3與式（5)中之 R 1或R2相同。） 具體地，其包含碳酸烷基苯基酯，例如碳酸甲基苯基 10 312/發明說明書(補件)/92-10/92120780 200408622 醋、碳酸乙基苯基酯、碳酸丙基苯基酯、碳酸丁基苯基酯、 石反酸己基苯基酯；以及碳酸曱基甲苯基酯、碳酸乙基曱苯 基醋、碳酸曱基二曱苯基酯、碳酸乙基二甲苯基酯。 本發明方法中之另一產物碳酸二芳酯係以下式（8)表 示：ArOH (6) (where Ar represents an aromatic group having 6 to 20 carbon atoms.) Specifically, it contains phenol, ortho-cresol, m-'cresol or p-cresol, ortho-ethyl Phenol, m-ethylphenol or p-ethylphenol, o-propylphenol, m-propylphenol or p-propylphenol, o-methoxyphenol, m-methoxyphenol or p-methoxy Phenol, 2, 6-dimethylphenol, 2, 4-dimethylphenol, 3, 4-dimethylphenol, o-chlorophenol, m-chlorophenol or p-chlorophenol, 1-naphthol, 2-naphthol. Particularly preferred is phenol. One of the products of the method of the present invention is alkylaryl carbonate which is represented by the following formula (7): R3-0C00-Ar (7) (wherein Ar has the same meaning as in formula (6); and R3 is in formula (5) R 1 or R 2 is the same.) Specifically, it contains an alkyl phenyl carbonate, such as methyl phenyl carbonate 10 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 Vinegar, ethyl phenyl carbonate Propyl carbonate, butyl phenyl carbonate, hexyl phenyl carbonate, fluorenyl tolyl carbonate, ethyl fluorenyl carbonate, fluorenyl diphenyl phenyl carbonate, ethyl carbonate Xylyl ester. The diaryl carbonate, another product in the method of the present invention, is represented by the following formula (8):
Ar-0C00-Ar (8) (其中Ar具有與式（6)中相同的意義。） 具體地，其包含碳酸二苯基酯、碳酸二—曱苯基酯、碳 ^ 一甲本基醋、$反酸二萘基g旨、碳酸雙（氯苯基）醋。 適用於本發明之觸媒可為任一種可促進碳酸二烷酯或 石反酸燒基芳酯與芳族羥基化合物之轉酯化反應及碳酸烷基 芳醋之不相稱反應者。舉例來說，此觸媒包含如下： (a) 錫化合物，例如 BU2sn〇、Ph2SnO、（C8H”）2SnO、Bu2Sn(〇Ph)2、 Bu2Sn(〇CH3)2 、 Bu2Sn(0Et)2 、 Bu2Sn(0Ph)0 (0Ph)SnBu2 ; (b) 錯化合物，例如 Pb〇、pb(〇ph)2、pb(〇c〇CH3)2 ; (c) 路易士酸化合物，例如 A1X3、ηχ3、TiX4、ZnX2、FeX3、SnX4、 VX5 (其中X代表鹵素原子、乙醯氧基基團、烷氧基基團或芳氧基 基團）’具體地為 A1C13、Al(0Ph)3、TiCl4、Ti(0Ph)4、Ti(0Et)4、 Ti(0Pr)4 、 Ti(0Bu)4 ; (d) 錯化合物，例如Zr (acac)4、Zr〇2(其中acac代表乙醯基絡 合物配位基）； (e) 銅化合物，例如 CuCl、C11CI2、CuBr、CuBr〗、Cul、Cul2、 Cu(OAc)2(其中Ac代表乙醯基基團。） 於上述化合物中，錫化合物及鈦化合物為特佳。 11 312/發明說明書(補件)/92-10/92120780 200408622 本發明之特徵在於其使用具至少二個串聯連接的獨立 反應區結構之反應裝置。個別反應區可為獨立的反應器或 可由藉分隔一反應器内部而形成之反應區段。其所必要的 條件為每一反應區中產生之蒸氣可獨立地取出。 圖1顯示具二個連接的反應器結構之反應裝置實施例。 於圖1之反應裝置中，反應器1係與反應器2相結合。反 應器1設有管線L 1 (高沸點起始材料自此處通過）、管線 L 2 (低沸點起始材料自此處通過）、管線L 4 (用以排出液體） 及管線V 1 (用以排出蒸氣）。反應器2設有上述的L 4、用以 排出液體之管線L 5及用以排出蒸氣之管線V 2。已通過管 線V 2排出之蒸氣加熱熱交換器1中之液體，以供能量回 收。必要時，反應器1中之反應液體可藉内部線圈或安裝 於反應器1之夾套加熱。 可獨立地控制反應器1及反應器2中之反應條件，因 此，可適當地設定此等反應器中之反應溫度及蒸氣產生， 因而可使自其之熱量回收達到最佳。舉例來說，當反應器 2中之溫度高於反應器1中溫度時，則可確保供熱交換所 需之溫度差，且反應器2中產生的蒸氣熱量可有效率地轉 移至反應器1中之液體。 於本發明具至少二個連接的反應器之反應裝置中，每一 反應器可設置用於熱回收之熱交換器。圖2顯示具三個連 接的反應器結構之反應裝置實施例。於此例中，低沸點起 始材料係經由三條管線L 2 a、L 2 b及L 2 c而分開地供給，且 藉每一反應器中產生的蒸氣之冷凝熱加熱，以致於部分地 12 312/發明說明書(補件)/92-10/92120780 200408622 氣化，並且生成的能量可用於加熱每一步驟中之反應液體。 於圖2之結構中（其中有三個連接的反應器1至3 )，所 有高沸點起始材料（管線L 1 )與部分低沸點起始材料（管線 L 2 a)混合，且供給至反應器1中。藉反應器1中產生的高 溫蒸氣，藉熱交換器1 (管線V 1 )之作用加熱生成的低溫混 合物，且反應器1中一部分必要的能量因而供給於其中。 於熱交換後，高溫蒸氣部分地冷凝，且經由管線L 6取出。 經由排放管線L 3將反應液體轉移至反應器2。 當使得反應器1中之溫度與反應器2之溫度相同時，通 過排放管線L 3供給之液體無法藉反應器2中產生之蒸氣加 熱。因此，反應器2中產生之高溫蒸氣加熱及氣化低沸點 起始材料（經由管線L 2 b分開地供給）。因此，氣化的低沸 點起始材料與反應器2中之反應液體接觸，且因而液化， 並且其冷凝熱提供能量予反應液體。於熱交換後，上述的 高沸點蒸氣部分地冷凝為液體。當含於蒸氣中之低沸點產 物返回反應器時，其可能妨礙轉酯化反應。然而，由於反 應器2中形成之低沸點產物相當地分布於蒸氣中，但少量 分布於冷凝的液體中，故冷凝的液體可能在蒸氣-液體分離 器1中分離，且可能經由管線L 7在反應器1中回收。含大 量低沸點產物之蒸氣係經由管線V4取出，且反應器2中之 反應液體係經由管線L4而轉移至反應器3。 於反應器3中，經由管線L 2 c而分開地供給之低沸點起 始材料可能同樣地在反應器2中回收高溫蒸氣之熱量，且 來自蒸氣之冷凝液體在蒸氣-液體分離器2中分離。經由管 13 312/發明說明書(補件)/92-10/92120780 200408622 線L8，在反應器2中回收此熱量，且經由管線V5取出系 統之蒸氣。經由管線L 5取出反應器3中之反應液體。 當4或多個反應器於此中連接時，可以上述相同的方式 連續地驅動反應器。 一般而言，塞流式反應器能提供高的材料轉化率（相對 於連續攪拌槽反應器）。然而，就需要連續地自反應系統取 出產物之反應器而言，規劃適合的塞流式反應器是困難 的。當連續攪拌槽反應器串聯連接時，反應液體可能流入 液相區域（如同於塞流式反應器中）。然而，具有安裝複數 個攪拌槽之工廠具缺點為，安裝其上之攪拌槽數目增加而 明顯地提高設備成本。 因此，為了保留塞流性質，且為了降低其設備成本，使 一反應器分隔為複數個區段且使用具因而分隔的複數個區 段反應器之反應裝置是必要的。 於此類型反應裝置之一實例中，反應器係藉著於其中設 置較低隔間與較高隔間之組合而分隔為複數個反應區段， 依此方式，較低隔間將反應器之較低部分分開為複數個密 閉的區段（除了其中液相流之通道以外），且其較高部分仍 於其上方具空間，且較高隔間將反應器之較高部分分開為 複數個供氣相用之密閉區段，且其較低部分仍於其下方具 空間。於反應器中，液相依序連續地自第一反應區穿過至 最終反應區，同時形成連續相，且藉其附近之液相使每一 反應區中之蒸氣相被密封起來，並且自反應裝置獨立地取 出蒸氣。可獨立地控制及固定每一反應區中除了壓力以外 14 312/發明說明劃補件)/92-10/92120780 200408622 之反應條件。因此，此類型反應器使得連續操作能以 2中相同方式徹底地進行。 以下將更詳細地說明反應器。 用於本發明之反應器具有特殊的結構。反應器較佳 其較低部分具「較低隔間」，係除了其中液相流之通玉 外，將液相區域分開為密閉區段，且其較高部分仍於 方具空間；且反應器於其較高部分具「較高隔間」，係 相區域分開為密閉區段，且其較低部分仍於其下方具 間。於具有隔間之反應器中，液相區域可能為至少二 開的區段，依此方式，其可能於附近區段間運轉，同 氣相可能為至少二個獨立的分開區段。此類型反應器 用於本發明中。於反應器中，一般而言，「較低隔間」 目與「較高隔間」之數目相同，且其係交替地配置。因 用此方式，反應器係分開為用於蒸氣相及液體相之「 區段」（其數目比較高或較低隔間數目大1個）。 舉例來說，圖5為顯示分隔為三個反應區段之反應 構實施例的示意斷面圖。於圖5中，反應器1設有分 中液相之較低隔間A 1及A 2以及分隔其中氣相之較高 B 1及B 2。於此等隔間中，較高隔間深深地浸入液相區 再者，反應器1具原料供應管線L1及L 2、蒸氣排放丨 L 3 a、L 3 b及L 3 c (自每一反應區段）以及反應液體排放 L4。雖然圖5中未顯示，但可能是必要的是於每一反 段設置内部線圈，或外部夾套可安裝於反應區周圍， 於此線圈及夾套中循環熱介質，藉以提供反應所需的 312/發明說明書(補件)/92-10/92120780 與圖 係於 I以 其上 將氣 空 個分 時蒸 較佳 之數 此， 反應 器結 隔其 隔間 域。 •線 管線 應區 且可 能量 15 200408622 至本發明反應裝置之每一反應區段中，或可於反應 安裝再沸器以提供熱予反應器。此效果之修飾通常 明中是較佳的。 如說明，於本發明使用之反應器中，依可於相鄰 運行之方式，藉較低隔間將液相區域分開為至少二 段，且反應器經設計使得液相可依序自第一區段運 後階段。用於使液相運行於相鄰區段間之裝置未特 定。舉例來說，液相可溢流通過每一較低隔間之頂苟 或其可流經每一較低隔間上方邊緣形成之凹槽；或 個孔形成於每一較低隔間之中間部分（液體可運行i 間）；或此等可以任一所欲方式合併。特別地，當僅 較低隔間頂部邊緣上方之溢流及/或藉通過每一較4 方邊緣間形成之凹槽達成相鄰隔間之間的液相流時 的是，每一較高隔間之較低邊緣可位於液相區域之 分（深度），俾控制反應液體之通道，以便特定液相 之反應液體保留時間分布可能不會太大。 一般而言，相較於連續攪拌槽反應器，塞流型反 能提供高的起始材料轉化率。然而，如上述，當反 設計使得液相可依序自所用之第一區段流至最終區 則反應液相部分可更容易地以類似塞流型反應器中 流動於其中（相較於事聯連接先前技藝已提出之連句 槽反應器）。 再者，本發明所用之反應器經設計為，含有副產 族醇）及碳酸二芳酯（於反應程序中製得）或其混合來 312/發明說明書(補件)/9240/92120780 器外部 在本發 區段間 個區 行至最 別地界 5邊緣； 一或多 I過其 藉每一 &隔上 ，必要 較低部 區段中 應為氣 應器經 段時， 之方式 ΐ攪拌 物（脂 7之輕 16 200408622 質德份可經由氣相排放管線L 3 a、L 3 b、L 3 c以氣相 排出。反應器中之氣相部分分開為獨立區段（由於其 封之液體），且經由每一區段之每一蒸氣排放管線取 同組成之蒸氣。於反應器之第一區段處，轉酯化程 當地高，因itb ，脂族醇之比例可能是大的；但相反 其較後階段中，轉酯化程度將是低的，因此，起始才 酸二芳酯及芳族羥基化合物）之比例可能是大的。因 其組成而定，可處理此等蒸氣，且可移除副產物（月 且可回收起始材料。相較於所有蒸氣在相同條件下 例子，此提高脂族醇分離過程之能量效率。為了藉 餾塔處理蒸氣（視其組成而定），一般而言，將個別 給至蒸餾塔之不同區域是必要的。甚至在攪拌槽串 之裝置中，可達成相同的蒸氣處理。然而，連接複 拌槽之裝置是昂貴的，且使用本發明反應器之方法 的。 此外，由於可提供熱予本發明反應器之每一區段 每一反應區段中之溫度可獨立地控制。因此，當較 之反應區段溫度高於較前階段中之溫度時，可促進 此外，較後階段中產生之蒸氣可用以加熱較前階段 應液體。此可進一步地提高本發明方法中之能量效 根據本發明方法，依序將液相自第一區段至最終 行，且連續地將觸媒及芳族羥基化合物引入反應器 分之第一區段中（當其於液相條件時），且連續地將 芳酯引入反應器之至少一反應區段（當其於氣相或浓 312/發明說明_ 補件)/92-10/92120780 連續地 周圍密 出具不 度將相 地，於 才料（碳 此，視 旨族醇） 處理之 使用蒸 蒸氣供 聯連接 數個攪 是經濟 中，故 後階段 反應， 中之反 率〇 區段運 液相部 碳酸二 L相條 17 200408622 件時）。同樣地，芳族羥基化合物可引入反應器之至少一反 應區段中（當其於氣相或液相條件時）。依此方式，將碳酸 二芳s旨或芳族經基化合物引入部分或全部分開的反應區段 (當其於氣相或液相條件時），因而可促進副產物（脂族醇） 之蒸發作用，且可有利地將反應平衡引至製造系統。 雖然附圖中未顯示，但可於每一反應區段設置内部線 圈，或外部夾套可安裝於反應區周圍，且可於此線圈及夾 套中循環熱介質，藉以提供反應所需的能量至本發明反應 裝置之每一反應區段中，或可於反應器外部安裝再沸器以 提供熱予反應器。此效果之修飾通常在本發明中是較佳的。 反應區（反應區段）之數目可為2或更多，且未特別地界 定。然而，即使數目較所需為高，效果將會逐漸地下降。 因此，數目通常可為2至30，較佳為2至15。 可藉外部攪拌器強迫地攪拌構成此中個別反應區之每 一反應器及反應器中之個別反應區，但此攪拌不是必要 的。自然流動或對流流動或冒泡（以蒸發作用可產生者）可 能足夠混合反應物質。此中可使用之外部攪拌系統包含例 如具攪拌葉片者、具泵送循環者及具蒸氣或蒸汽吹動者。 於本發明製造芳族碳酸酯之方法中，於氣相或液相中之 碳酸二芳酸或芳族羥基化合物係引入反應器中之部分或全 部分隔的反應區中，同時促進副產物（脂族醇）之蒸發作 用，藉以有利地將反應平衡引導至製造系統中。此外，大 部分在較後階段形成之脂族醇係自系統取出，不與前階段 中之反應液體混合，因此，脂族醇不會累積於較前階段之 18 312/發明說明書(補件)/92-10/92120780 200408622 反應區中（不會妨礙反應進行）。 本發明之方法通常不需溶劑，但可使用對於反應呈惰性 之溶劑（例如選自醚類、脂族烴及芳族烴）。於反應裝置（反 應器）中最終反應區（反應區段）取出之液相受到純化構件 (例如通過蒸餾作用）作用，且可製得所欲之碳酸烷基芳酯 及/或碳酸二芳酯。 雖然視所用之起始材料類型及組成而定，本發明方法中 之反應溫度通常可落於5 0及3 0 0 °C間，較佳介於1 0 0及 2 5 0 °C間。於較高溫度下，反應速率將較高，但副產物（烷 基芳族醚等）將於高溫增加。因此，太高溫度不利於本發明 之方法。視反應器中之起始材料及組成物類型而定，可改 變反應器内之壓力。一般而言，可於壓力下或減壓下（介於 10及3000kPa間，尤其是介於50及2000kPa間（介於0.5 及2 0大氣壓間））進行。 一般而言，觸媒係溶解或分散於起始材料中，且供給至 反應區（反應器）中。以欲供給的起始材料為基礎，觸媒含 量通常為0 . 0 0 0 1至1 0莫耳％，較佳為0 . 0 0 1至5莫耳％。 倘若觸媒含量太低時，反應速率將是不足的；但倘若太低 時，可能提高副產物（烷基芳族醚）含量。視所用之其他反 應條件而定，可改變反應裝置（反應器）中液體之平均滯留 時間，但通常落於0. 1及2 0小時之間，較佳介於0. 3及 1 0小時之間。 〈實施例〉 本發明之具體例將參照以下實施例更進一步地說明，然 19 312/發明說明書(補件)/92-10/92120780 200408622 而，本發明未脫離其精神而受限於此等具體例。 [實施例1 ] 本實施例使用二個反應器，其中後階段之反應溫 前階段之反應溫度，且前階段中之反應液體係藉後 生供有效率能量利用之高溫蒸氣加熱。 於圖1所示其中具二個3 0 0毫升反應器之反應系 以酚使碳酸二甲酯轉酯化。每一反應器中之液體液位 為約50% 。 經由液體進料管線L 1，供給9 4克/小時（1莫耳/小 酚及0 . 6克/小時觸媒（氧化二丁基錫），並且經由液體 管線L 1，供給9 0克/小時（1莫耳/小時）碳酸二甲酉旨。 的起始材料之溫度為6 0 °C。於二反應器中壓力為5 0 0 且反應器1中之溫度為180 °C ，反應器2中之溫度為 〇C。 反應液體係藉反應器2中產生之蒸氣在熱交換器1 熱，且接著依序引導至管線L3、反應器1、管線L4及 器2，且最後經由液體排放管線L 5取出。經由蒸氣排 線V1取出反應器1中產生之蒸氣。經由管線V 2，將 器2中產生之蒸氣引導至熱交換器1，且當於其中冷；I 使用其加熱起始材料液體，且接著經由管線L 6排出 於此條件下，使裝置連續驅動8小時，且收集通過 L 5排出之反應液體及分析其組成。於此分析中，可偵 5 . 9重量％碳酸甲基苯酯及1 . 6重量％碳酸二苯基酯。 管線L 5之流率為1 2 4克/小時，且產物碳酸甲基苯酯 312/發明說明書(補件)/92] 0/92120780 .高於 •段產 丨中， 保持 時） 進料 供給 kPa， 200 中加 反應 放管 反應 "夺， 〇 管線 測到 通過 及碳 20 200408622 酸二苯基酯相當於總共0 . 0 5 7莫耳/小時。用於反應所需之 熱量為4 9 . 6千焦耳/小時。 每莫耳產物之能量效率為Π.5χ1(Γ4莫耳/千焦耳。 [實施例2 ] 本實施例使用三個反應器，其中每一階段中之起始材料 液體係每一階段產生供有效率能量利用之高溫蒸氣加熱。 於圖2所示其中具三個2 0 0毫升反應器之反應系統中， 以酚使碳酸二甲酯轉酯化。每一反應器中之液體液位保持 為約50°/〇。 經由液體進料管線L 1，連續地供給與實施例1相同量之 酚及觸媒（氧化二丁基錫），並且經由液體進料管線L 2且分 開地經由管線L 2 a、L 2 b及L 2 c，供給與實施例1相同量之 碳酸二曱酯。將通過管線L 2 a、L 2 b及L 2 c之流率控制為比 例 8 / 1 / 1。 關於其條件，全部三個反應器係在2 0 0 °C及在5 0 0 k P a。 依序將反應液體引導至反應器1、管線L 3、反應器2、管 線L 4及反應器3，且最後經由液體排放管線L 5排出。 將來自管線L1之酚及觸媒與來自管線L 2 a之碳酸二甲 酯混合，且接著引導至熱交換器1。於熱交換器1中，藉 反應器1中產生之蒸氣加熱起始材料液體，且接著引導至 反應器1中。於用於熱交換後，於8 0 °C使蒸氣冷凝為液體， 且經由管線L 6將此排出。 經由管線V 2，將反應器2中產生之2 0 0 °C蒸氣引導至熱 交換器2，且將此來自管線L 2 b之受熱的碳酸二甲酯接著 21 312/發明說明書(補件)/92· 10/9212〇780 200408622 冷卻至1 6 6 °C 。於蒸氣-液體分離器1中分離生成的冷凝 液，且使其經由管線L7返回反應器1，同時使未冷凝的蒸 氣經由管線V4排出。於熱交換器2中加熱來自管線L2b 之受熱的碳酸二甲酯，且將1 5 8 °C之蒸氣供給至反應器2 中 〇 經由管線V 3，將反應器3中產生之2 0 0 °C蒸氣引導至熱 交換器3 (類似反應器之蒸氣），且將此來自管線L 2 c之受 熱的碳酸二曱酯接著冷卻至1 5 0 °C。於蒸氣-液體分離器2 中分離生成的冷凝液，且使其經由管線L 8返回反應器2， 同時使未冷凝的蒸氣經由管線V 5排出。於熱交換器3中加 熱來自管線L 2 c之受熱的碳酸二甲酯，且將1 5 8 °C之蒸氣 供給至反應器3中。 於此條件下使裝置連續驅動8小時，且收集通過管線L 5 排出之反應液體及分析其組成。於此分析中，可偵測到6. 4 重量％碳酸曱基苯酯及1 . 8重量％碳酸二苯基酯。通過管線 L 5之流率為1 2 7克/小時，且產物碳酸甲基苯酯及碳酸二 苯基酯相當於總共0 . 0 6 5莫耳/小時。用於反應所需之熱量 為5 3 . 3千焦耳/小時。 每莫耳產物之能量效率為12. 2xl(T4莫耳/千焦耳，且其 係高於實施例1中者。 [比較例1 ] 於本例中，反應液體係與如實施例2中相同的方法運 轉，但未進行自產生蒸氣之熱回收。 於圖3所示其中具三個2 0 0毫升反應器之反應系統中， 22 312/發明說明書(補件)/92-10/92120780 200408622 以酚使碳酸二甲酯轉酯化。每一反應器中之 為約50% 。 經由圖3中之液體進料管線L1，連續地供 相同量之酚及觸媒（氧化二丁基錫），並且經 線L 2且分開地經由管線L 2 a、L 2 b及L 2 c，供 相同量之碳酸二曱酯。將通過管線L 2 a、L 2 b 控制為比例5/2/3 。 關於其條件，全部三個反應器係在2 0 0 °C , 依序將反應液體引導至反應器1、管線L 3、 線L4及反應器3，且最後經由液體排放管線 由蒸氣排放管線V1，將反應器1中產生之蒸 管線V2，將反應器2中產生之蒸氣引導至冷 至1 6 0 °C ，且經由管線L 6將生成的冷凝液返 中，同時使未冷凝的蒸氣經由蒸氣排放管線 管線V 3，反應器3中產生之蒸氣亦以如上述 冷凝器中，且冷卻至1 6 0 °C，且經由管線L 7 液返回反應器2中，同時使未冷凝的蒸氣經 線V5排出。 於此條件下使裝置連續驅動8小時，且收； 排出之反應液體及分析其組成。於此分析中， 重量％碳酸甲基苯酯及1 . 8重量％碳酸二苯基 L 5之流率為1 3 5克/小時，且產物碳酸甲基J 苯基酯相當於總共0 . 0 7 2莫耳/小時。用於反 為9 3 . 4千焦耳/小時。 312/發明說明書(補件)/92-10/92120780 液體液位保持 r給與實施例1 由液體進料管 :給與實施例1 及L2c之流率 及在 500kPao 反應器2、管 L 5排出。經 氣排出。經由 凝器，且冷卻 回反應器1 V 4排出。經由 之方式引導至 將生成的冷凝 由蒸氣排放管 |通過管線L5 可偵測到6. 7 酯。通過管線 良酯及碳酸二 應所需之熱量 23 200408622 每莫耳產物之能量效率為7. 7xl(T4莫耳/千焦耳，且其 證實此低於先前實施例中者。 [比較例2 ] 本例使用三個串聯連接之反應器，其中起始材料酚與碳 酸二甲酯係彼此逆流地接觸。 於圖4所示其中具三個串聯連接的200毫升壓力鍋之反 應系統中，以彼此逆流接觸方式，以酚轉酯化碳酸二甲酯。 每一反應器中之液體液位保持為約5 0 %。 經由圖4中之液體進料管線L1，連續地供給與實施例1 相同量之酚及觸媒（氧化二丁基錫），並且經由液體進料管 線L 2，供給與實施例1相同量之碳酸二甲酯，其中此完全 地氣化，且接著經由管線V 1 a供給至反應器3。 反應器中之壓力經控制使得可自反應器1朝反應器3提 高，以供蒸氣流動。具體地，反應器1中之壓力為5 0 0 k P a， 反應器2中之壓力為550kPa，且反應器3中之壓力為600 k P a。關於反應溫度，反應器3經加熱使得其溫度可為2 0 0 °C ，但反應器1及反應器2未加熱。 依序將反應液體通過反應器1、管線L 3 a、反應器2、管 線L 3 b及反應器3，且最後經由液體排放管線L 3 c排出。 經由管線V 1 b，將反應器3中產生之2 0 0 °C蒸氣供給至 反應器2 ;經由管線V 1 c，將反應器2中產生之蒸氣供給至 反應器1 ;且經由管線V 2，將反應器1中產生之蒸氣排出。 於此條件下使裝置連續驅動8小時，且收集通過管線L 3 c 排出之反應液體及分析其組成。於此分析中，可偵測到7 . 4 24 312/發明說明書(補件)/92-10/92120780 200408622 重量％碳酸甲基苯酯及1 . 6重量％碳酸二苯基酯。通過管線 L 3 c之流率為1 4 9克/小時，且產物碳酸甲基苯酯及碳酸二 苯基酯相當於總共0 . 0 8 4莫耳/小時。用於反應所需之熱量 為7 3 . 2千焦耳/小時。 每莫耳產物之能量效率為11. 5x10 _4莫耳/千焦耳，且其 證實其類似實施例1中者，但低於實施例2中者。 如實施例1及2以及比較例1及2，根據本發明之方法， 液相部分通過分開的複數個串聯反應區，而輕質餾份經蒸 除，因此輕質餾份（例如可能造成對於起始系統反應平衡之 脂族醇）可能降至最小程度，而不需任何額外及明顯的設備 成本，此外，可容易地確保充足的反應體積。因此，根據 本發明之方法使得連續製造芳族聚碳酸酯成為可能，且其 生產率及選擇率同樣地是高的。此外，於本發明之方法中， 可於廣泛範圍内容易地選擇反應溫度、反應壓力、滯留時 間及其他條件，且本發明之又另一優點為可使用極簡化的 裝置以達成極高生產率及選擇率之產物。 [實施例3 ] 使用分隔為三個反應區段（如圖5之結構）之6 0 0毫升反 應器，根據圖6所示之反應系統，以酚使碳酸二甲酯轉酯 化。 如圖6，經由液體進料管線L1，連續地供給9 4克/小時 (1莫耳/小時）酚及〇 . 6克/小時觸媒（氧化二丁基錫），並 且經由液體進料管線L 2且分開地經由管線L 2 a、L 2 b及 L 2 c，供給9 0克/小時（1莫耳/小時）之碳酸二曱S旨至每一 25 312/發明說明書(補件)/92-10/92120780 200408622 反應區段。將通過管線L 2 a、L 2 b及L 2 c之流率控制為 5/2/3 ° 反應器中之壓力為500 kPa，且其中溫度為200 °C 反應液體自第一區段溢流至第二區段及至第三區段， 著經由液體排放管線L 3排出。經由蒸氣排放管線V1 第一區段中產生之蒸氣排出。經由管線V 2，將第二區 產生之蒸氣引導至冷凝器，且冷卻至1 6 0 °C ，且使生 冷凝液經由管線L4返回第一區段，同時使未冷凝的蒸 由蒸氣排放管線V 4排出。經由管線V 3，亦將第三區 產生之蒸氣引導至冷凝器，且冷卻至1 6 0 °C ，且使生 冷凝液經由管線L 5返回第二區段，同時使未冷凝的蒸 由蒸氣排放管線V 5排出。 於此條件下使裝置連續驅動8小時，且收集通過管 排出之反應液體及分析其組成。於此分析中，可偵測ί丨 重量％碳酸曱基苯酯及1 . 8重量％碳酸二苯基酯。通過 L 3之流率為1 3 5克/小時，且產物碳酸甲基苯酯及碳 苯基酯相當於總共0 . 0 7 2莫耳/小時。用於反應所需之 為9 3 . 4千焦耳/小時。反應器中之液體液位保持為約 每莫耳產物之能量效率為7. 7χ1(Γ4莫耳/千焦耳。 [比較例3 ] 連續地供給9 4克/小時（1莫耳/小時）酚、0 · 6克/ λ 觸媒（氧化二丁基錫）及9 0克/小時（1莫耳/小時）碳酸 酯至5 0 0毫升壓力鍋中，同時使反應器中之液體液位 為約6 0 %。反應器中之壓力為5 0 0 k P a，且其中溫度為 312/發明說明書(補件)/92_ 10/9212〇780 比例 〇使 且接 ，將 段中 成的 氣經 段中 成的 氣經 線L3 16.7 管線 酸二 熱量 5 0% ° 丨、時 二曱 保持 200 26 200408622 °c 。連續地排出反應器中產生之蒸氣。 連續進行反應8小時，且收集反應液體及分析其組成。 於此分析中，可偵測到5. 9重量％碳酸甲基苯酯及1 · 6重量 %碳酸二苯基酯。反應液體之排放速率為1 2 2克/小時。產 物碳酸甲基苯酯及碳酸二苯基酯相當於總共0 . 0 5 7莫耳/ 小時。用於反應所需之熱量為8 2. 0千焦耳/小時。 每莫耳產物之能量效率為7. Οχ 1(Γ4莫耳/千焦耳，且每 能量之生產率小於實施例3中者。 [比較例4 ] 使用圖7之反應器，其中液相分隔為三個區段，但沒有 較高隔牆（蒸氣相受此分隔），因此，蒸氣相形成連續蒸氣 相，以酚轉酯化碳酸二甲酯。 具體地，經由液體進料管線L1，供給與實施例3相同量 之酚及觸媒（氧化二丁基錫），並且經由液體進料管線 L 2 a、L 2 b及L 2 c，供給與實施例3相同量之碳酸二甲酯至 個別反應區段。然而，將通過管線L 2 a、L 2 b及L 2 c之流率 控制為比例6/2/2。 反應器中之壓力為500 kPa，且其中溫度為200 °C 。使 反應液體自第一區段溢流至第二區段及至第三區段，且接 著經由液體排放管線L 3排出。於每一反應區段中產生之蒸 氣在反應器中變得均勻，且經由蒸氣排放管線V1排出。 於此條件下使裝置連續驅動8小時，且收集通過管線L3 排出之反應液體及分析其組成。於此分析中，可偵測到6 . 5 重量％碳酸甲基苯酯及1 . 8重量％碳酸二苯基酯。通過管線 27 312/發明說明書(補件)/92-10/92120780 200408622 L 3之流率為1 2 2克/小時，且產物碳酸曱基苯酯及碳酸二 苯基酯相當於總共0 . 0 6 2莫耳/小時。用於反應所需之熱量 為8 1 . 9千焦耳/小時。 每莫耳產物之能量效率為7. 6xl0_4莫耳/千焦耳。每能 量之生產率大於比較例1中者，但仍小於實施例3中者。 [實施例4 ] 使用圖5之反應器，根據圖8所示之反應系統，以酚使 碳酸二甲酯轉酯化。 如圖8，經由液體進料管線L1，連續地供給與實施例3 相同量之酚及觸媒（氧化二丁基錫），並且經由液體進料管 線L 2且分開地經由管線L 2 a、L 2 b及L 2 c，供給碳酸二甲 酯。將通過管線L 2 a、L 2 b及L 2 c之流率控制為比例8 / 1 / ：1。 使來自管線L 1之酚及觸媒與來自管線L 2 a之碳酸二甲 酯混合，且經由管線L 3供給至熱交換器E 1，藉反應器第 一區段中產生之蒸氣加熱起始材料，且接著進入反應器第 一區段。於歷經熱交換後，蒸氣於8 0 °C變為冷凝液，且經 由管線L 6排出。 反應器中之壓力為500 kPa，且其中溫度為200 °C 。使 反應液體自第一區段溢流至第二區段及至第三區段，且接 著經由液體排放管線L 5排出。 經由管線V 2，將第二區段中產生之2 0 0 °C蒸氣引入熱交 換器E 2，接著加熱來自管線L 2 b之碳酸二甲酯，且冷卻至 1 6 6 °C 。於蒸氣-液體分離器中分離生成的冷凝液，且使其 經由管線L4a返回第一區段，同時使未冷凝的蒸氣經由管 28 312/發明說明書(補件)/92-10/92120780 200408622 線V 4排出。於熱交換器E 2中加熱來自管線L 2 b之 碳酸二甲酯，且將1 5 8 °C之蒸氣供給至反應器之第 中 〇 經由管線V3，將第三區段中產生之2 0 0 °C蒸氣引 交換器E 3 (類似第二區段中之蒸氣），且將此來自管 之受熱的碳酸二曱酯接著冷卻至1 5 0 °C。於蒸氣-液 器中分離生成的冷凝液，且使其經由管線L 4 b返回 段，同時使未冷凝的蒸氣經由管線V 5排出。於熱ί Ε 3中加熱來自管線L 2 c之受熱的碳酸二甲S旨，且將 之蒸氣供給至第三區段中。 於此條件下使裝置連續驅動8小時，且收集通過 排出之反應液體及分析其組成。於此分析中，可偵測 重量％碳酸曱基苯酯及1 . 8重量％碳酸二苯基酯。通 L 5之流率為1 2 7克/小時，且產物碳酸甲基苯酯及 苯基酯相當於總共0 . 0 6 5莫耳/小時。用於反應所需 為5 3 . 3千焦耳/小時。 每莫耳產物之能量效率為12. 2 Οχ 1(Γ4莫耳/千焦 是高的。 [比較例5 ] 使三個2 0 0毫升壓力鍋串聯連接，並且根據圖9 系統，以彼此逆流接觸方式，以酚轉酯化碳酸二曱 中起始材料酚與碳酸二甲酯係彼此逆流地接觸。 如圖9，經由液體進料管線L1，連續地供給與實 相同量之酚及觸媒（氧化二丁基錫）至反應器R 1。經 312/發明說明書(補件)/92-10/92120780 受熱的 二區段 導至熱 .線 L 2 c 體分離 第二區 t換器 1 58〇C 管線L5 到6. 4 過管線 碳酸二 之熱量 耳，且 之反應 酯。其 施例1 由液體 29 200408622 進料管線L 2，供給碳酸二甲酯至熱交換器E 1，其中此 地氣化，且接著經由管線V 1 a供給至反應器R 3。 反應器中之壓力經控制使得可自反應器R 1朝反應i 提高，以供蒸氣流動。具體地，反應器R1中之壓力為 kPa，反應器R2中之壓力為550kPa，且反應器R3中 力為6 0 0 k P a。關於反應溫度，反應器R 3經加熱使得 度可為200 °C ，但反應器R1及反應器R2未加熱。 依序將反應液體通過反應器R 1、管線L 3 a、反應器 管線L 3 b及反應器R 3，且最後經由液體排放管線L 3 c去 經由管線V 1 b，將反應器R 3中產生之2 0 0 °C蒸氣供 反應器2 ;經由管線V1 c，將反應器R 2中產生之蒸氣 至反應器R1 ;且經由管線V 2，將反應器R1中產生之 排出。 於此條件下使裝置連續驅動8小時，且收集通過管綠 排出之反應液體及分析其組成。於此分析中，可偵測到 重量％碳酸甲基苯酯及1 . 6重量％碳酸二苯基酯。通過 L 3 c之流率為1 4 9克/小時，且產物碳酸甲基苯酯及碳 苯基酯相當於總共0 . 0 8 4莫耳/小時。用於反應所需之 為7 3 . 2千焦耳/小時。 每莫耳產物之能量效率為11.47xl(T4莫耳/千焦耳 是高的，但低於實施例4中者。 如實施例3及4以及比較例3、4及5，本發明之特 應器使得透過芳族羥基化合物與碳酸二烷酯之轉酯化 連續製造芳族碳酸酯成為可能，且其能量效率是高的 312/發明說明書(補件)/92-10/92120780 完全 i R3 500 之壓 其溫 R2、 卞出。 給至 供給 蒸氣 .L3c 7.4 管線 酸二 熱量 ，且 殊反 反應 30 200408622 般而言，轉酯化反應是平衡反應，且大大地轉移至起始系 統，且其反應速率是低的。 本發明已參照其特殊具體例而詳細地說明，但熟習本技 藝之人士當可明白，可進行許多變化及修飾，而不脫離其 精神及範圍。 本申請案係以2 0 0 2年8月1 2申請之曰本專利申請案（申 請案第2 0 0 2 - 2 3 4 8 8 4號）及2 0 0 2年8月1 3中請之日本專利 申請案（申請案第2 0 0 2 - 2 3 5 3 8 5號）為基礎，且其内容係合 併於本案以供參考。 〈工業應用性〉 根據本發明之方法，液相部分通過分開的複數個串聯反 應區，而輕質餾份經蒸除，因此輕質餾份（例如可能造成對 於起始系統反應平衡之脂族醇）可能降至最小程度，而不需 任何額外及明顯的設備成本，此外，可容易地確保充足的 反應體積。因此，根據本發明之方法使得連續製造芳族聚 碳酸酯成為可能，且其生產率及選擇率同樣地是高的。此 外，於本發明之方法中，可於廣泛範圍内容易地選擇反應 溫度、反應壓力、滯留時間及其他條件，且本發明之又另 一優點為可使用極簡化的裝置以達成極高生產率及選擇率 之產物。 【圖式簡單說明】 圖1為顯示具二個反應器串聯連接結構之反應裝置實施 例的示意斷面圖。 圖2為顯示具三個反應器串聯連接結構之反應裝置實施 31 312/發明說明書(補件)/92-10/92120780 200408622 例的示意斷面圖。 圖3為顯示比較例1之反應器結構及反應系統的示意斷 面圖。 圖4為顯示比較例2之反應器結構及反應系統的示意斷 面圖。 圖5為顯示内部分隔為三個反應區之反應器結構實施例 的示意斷面圖。 圖6為顯示實施例3之反應器結構及反應系統的示意斷 面圖。 圖7為顯示比較例4之反應器結構及反應系統的示意斷 面圖。 圖8為顯示實施例4之反應器結構及反應系統的示意斷 面圖。 圖9為顯示比較例5之反應器結構及反應系統的示意斷 面圖。 (元件符號說明） 1、2、3、R 1、R 2、R 3 反應器（反應槽） L 1、L 2 起始材料進料管線 L 3、L 4、L 5、L 6、L 7、L 8 液體排放管線 V 1、V 2、V 3、V 4、V 5 蒸氣排放管線 E 1、E 2、E 3 熱交換器 A 1、A 2 較低隔間Ar-0C00-Ar (8) (wherein Ar has the same meaning as in formula (6).) Specifically, it contains diphenyl carbonate, di-fluorenyl carbonate, carbon ^ methylbenzyl vinegar, $ Inverse acid dinaphthyl g, bis (chlorophenyl) carbonate. The catalyst suitable for the present invention may be any one which can promote the transesterification reaction of dialkyl carbonate or rosal acid aryl ester with aromatic hydroxy compounds and the disproportionate reaction of alkyl carbonate aromatic vinegar. For example, the catalyst includes the following: (a) Tin compounds, such as BU2sn0, Ph2SnO, (C8H ") 2SnO, Bu2Sn (〇Ph) 2, Bu2Sn (〇CH3) 2, Bu2Sn (0Et) 2, Bu2Sn ( 0Ph) 0 (0Ph) SnBu2; (b) wrong compounds, such as Pb0, pb (〇ph) 2, pb (〇c〇CH3) 2; (c) Lewis acid compounds, such as A1X3, ηχ3, TiX4, ZnX2 , FeX3, SnX4, VX5 (where X represents a halogen atom, an ethoxy group, an alkoxy group, or an aryloxy group) 'is specifically A1C13, Al (0Ph) 3, TiCl4, Ti (0Ph) 4. Ti (0Et) 4, Ti (0Pr) 4, Ti (0Bu) 4; (d) Wrong compounds, such as Zr (acac) 4, Zr〇2 (where acac stands for ethenyl complex ligand) (E) Copper compounds, such as CuCl, C11CI2, CuBr, CuBr, Cul, Cul2, Cu (OAc) 2 (where Ac represents an ethanoyl group.) Among the above compounds, tin compounds and titanium compounds are particularly preferred. 11 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 The present invention is characterized in that it uses a reaction device having at least two independent reaction zone structures connected in series. The individual reaction zones may be independent reactors or may be Borrow points The reaction zone formed inside a reactor. The necessary condition is that the vapor generated in each reaction zone can be taken out independently. Figure 1 shows an example of a reaction device with two connected reactor structures. In the reaction device, the reactor 1 is combined with the reactor 2. The reactor 1 is provided with a line L 1 (a high boiling point starting material passes here) and a line L 2 (a low boiling point starting material passes here) Line L 4 (for discharging liquid) and Line V 1 (for discharging steam). Reactor 2 is provided with L 4 described above, line L 5 for discharging liquid and line V 2 for discharging vapor. The steam discharged through the line V 2 heats the liquid in the heat exchanger 1 for energy recovery. If necessary, the reaction liquid in the reactor 1 can be heated by an internal coil or a jacket installed in the reactor 1. Can be independently controlled The reaction conditions in Reactor 1 and Reactor 2, so the reaction temperature and steam generation in these reactors can be set appropriately, so that the heat recovery from them can be optimized. For example, when reactor 2 Medium temperature is higher than that of reactor 1 Temperature, it can ensure the temperature difference required for heat exchange, and the vapor heat generated in the reactor 2 can be efficiently transferred to the liquid in the reactor 1. The reaction in the present invention has at least two connected reactors In the device, each reactor may be provided with a heat exchanger for heat recovery. Fig. 2 shows an embodiment of a reaction device having three connected reactor structures. In this example, the low-boiling starting materials are separately supplied via three lines L 2 a, L 2 b, and L 2 c, and are heated by the condensation heat of the steam generated in each reactor, so that partly 12 312 / Invention (Supplement) / 92-10 / 92120780 200408622 Gasification, and the energy generated can be used to heat the reaction liquid in each step. In the structure of FIG. 2 (in which there are three connected reactors 1 to 3), all the high-boiling starting materials (line L 1) are mixed with part of the low-boiling starting materials (line L 2 a) and supplied to the reactor 1 in. The high-temperature steam generated in the reactor 1 is used to heat the low-temperature mixture generated by the action of the heat exchanger 1 (line V 1), and a part of the necessary energy in the reactor 1 is supplied thereto. After the heat exchange, the high-temperature vapor is partially condensed and taken out via line L 6. The reaction liquid is transferred to the reactor 2 via a discharge line L 3. When the temperature in the reactor 1 is made the same as the temperature in the reactor 2, the liquid supplied through the discharge line L3 cannot be heated by the vapor generated in the reactor 2. Therefore, the high-temperature steam generated in the reactor 2 heats and vaporizes the low-boiling-point starting material (separately supplied via the line L 2b). Therefore, the vaporized low-boiling point starting material comes into contact with the reaction liquid in the reactor 2, and thus liquefies, and its heat of condensation provides energy to the reaction liquid. After the heat exchange, the above-mentioned high-boiling-point vapor is partially condensed into a liquid. When the low-boiling products contained in the vapor are returned to the reactor, they may hinder the transesterification reaction. However, since the low-boiling products formed in the reactor 2 are fairly distributed in the vapor, but in a small amount in the condensed liquid, the condensed liquid may be separated in the vapor-liquid separator 1 and may be passed through the line L 7 in Recycled in reactor 1. The vapor containing a large amount of low-boiling products was taken out through line V4, and the reaction liquid system in reactor 2 was transferred to reactor 3 through line L4. In the reactor 3, the low-boiling starting material separately supplied via the line L 2 c may similarly recover the heat of the high-temperature vapor in the reactor 2, and the condensed liquid from the vapor is separated in the vapor-liquid separator 2 . This heat is recovered in reactor 2 via pipe 13 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 line L8, and the system vapor is removed via line V5. The reaction liquid in the reactor 3 is taken out through the line L 5. When 4 or more reactors are connected here, the reactors can be continuously driven in the same manner as described above. In general, plug flow reactors provide high material conversion rates (relative to continuous stirred tank reactors). However, it is difficult to plan a suitable plug flow reactor for a reactor that needs to continuously take out products from the reaction system. When a continuous stirred tank reactor is connected in series, the reaction liquid may flow into the liquid phase region (as in a plug flow reactor). However, a factory having a plurality of mixing tanks has a disadvantage in that the number of mixing tanks installed thereon increases the equipment cost significantly. Therefore, in order to retain the plug flow property and to reduce the equipment cost thereof, it is necessary to separate a reactor into a plurality of sections and use a reaction device having a plurality of zone reactors thus separated. In one example of this type of reaction device, the reactor is divided into a plurality of reaction sections by providing a combination of a lower compartment and a higher compartment therein. In this way, the lower compartment divides the reactor The lower part is divided into a plurality of closed sections (except for the channel in which the liquid phase flows), and the higher part still has space above it, and the higher compartment divides the higher part of the reactor into a plurality of Closed section for gas phase, and its lower part still has space below it. In the reactor, the liquid phase passes sequentially from the first reaction zone to the final reaction zone in succession, and a continuous phase is formed at the same time. The vapor phase in each reaction zone is sealed by the nearby liquid phase, and The reaction device takes out the vapor independently. Can independently control and fix the reaction conditions in each reaction zone except pressure 14 312 / Invention Note) / 92-10 / 92120780 200408622. Therefore, this type of reactor enables continuous operation to be performed thoroughly in the same manner as in 2. The reactor will be explained in more detail below. The reactor used in the present invention has a special structure. The lower part of the reactor preferably has a "lower compartment", except that the liquid phase area is divided into closed sections except for the flow of liquid phase, and the higher part is still in the square space; and The device has a "higher compartment" in its upper part, the phase area is divided into closed sections, and its lower part still has a space below it. In a reactor with compartments, the liquid phase region may be at least two separate sections. In this way, it may operate between nearby sections, and the co-phase may be at least two separate separate sections. This type of reactor is used in the present invention. In the reactor, in general, the number of "lower compartments" is the same as that of "higher compartments" and they are alternately arranged. In this way, the reactor is divided into "segments" for the vapor and liquid phases (the number of which is higher than the number of lower or higher compartments by one). For example, Fig. 5 is a schematic sectional view showing an example of a reaction structure divided into three reaction sections. In Fig. 5, the reactor 1 is provided with lower compartments A 1 and A 2 which separate the middle and liquid phases, and higher B 1 and B 2 which separate the gas phase therein. In these compartments, the higher compartment is deeply immersed in the liquid phase zone. Furthermore, the reactor 1 has a raw material supply line L1 and L2, and the steam is discharged. L 3 a, L 3 b, and L 3 c (from each A reaction section) and the reaction liquid discharge L4. Although it is not shown in Figure 5, it may be necessary to provide an internal coil in each reverse section, or an external jacket may be installed around the reaction zone, and a heat medium is circulated in this coil and jacket to provide the required reaction 312 / Invention Specification (Supplements) / 92-10 / 92120780 and Figures are based on I, which is the better number for time-steaming air and space, and the reactor is separated from its compartment area. • Lines, pipelines, and zones can be energized to each reaction zone of the reaction device of the present invention, or a reboiler can be installed in the reaction to provide heat to the reactor. Modifications of this effect are usually preferred. As illustrated, in the reactor used in the present invention, the liquid phase region is divided into at least two stages by a lower compartment in a manner that can be operated adjacently, and the reactor is designed so that the liquid phase can be sequentially from the first Post-transportation stage. The device used to run the liquid phase between adjacent sections is not specific. For example, the liquid phase can overflow through the top of each lower compartment or it can flow through the groove formed at the upper edge of each lower compartment; or a hole is formed in the middle of each lower compartment Portions (liquids can run between units); or these can be combined in any desired way. In particular, when only the overflow above the top edge of the lower compartments and / or the liquid phase flow between adjacent compartments is achieved by the grooves formed between each of the more square edges, each higher The lower edge of the compartment can be located in the liquid phase region (depth), and the channel of the reaction liquid is controlled so that the reaction liquid retention time distribution of a specific liquid phase may not be too large. In general, the plug flow type reaction provides a high conversion of the starting material compared to a continuous stirred tank reactor. However, as mentioned above, when the reverse design allows the liquid phase to flow sequentially from the first section to the final section, the reaction liquid phase section can more easily flow in a plug-flow-like reactor (as compared to the event (Lianlian Slot Reactor has been proposed in the previous art). In addition, the reactor used in the present invention is designed to contain by-product alcohols) and diaryl carbonate (made in the reaction process) or a mixture thereof. 312 / Invention Specification (Supplement) / 9240/92120780 outside the reactor Go between the sections of the hair section to the edge of the most special boundary 5; One or more I borrow each & partition, it is necessary to stir the way in the lower section when the reactor warp section, Matter (fat 7 of light 16 200408622) can be discharged in the gas phase through the gas phase exhaust lines L 3 a, L 3 b, L 3 c. The gas phase part in the reactor is divided into independent sections (due to its seal Liquid), and take the same composition of vapor through each vapor discharge line of each section. At the first section of the reactor, the transesterification process is locally high. Due to itb, the proportion of aliphatic alcohol may be large. But on the contrary, in the later stage, the degree of transesterification will be low, so the proportion of the starting diaryl ester and the aromatic hydroxy compound) may be large. Depending on its composition, these vapors can be processed and by-products can be removed (recyclable and starting materials can be recovered. This improves the energy efficiency of the aliphatic alcohol separation process compared to all vapors under the same conditions example.) Distillation towers are used to process steam (depending on their composition). Generally speaking, it is necessary to separately feed to different areas of the distillation tower. The same steam treatment can be achieved even in the device of the stirring tank string. However, the connection is complicated. The equipment of the mixing tank is expensive and uses the method of the reactor of the present invention. In addition, since the temperature in each reaction section of each section of the reactor of the present invention can be provided with heat, the temperature can be controlled independently. Therefore, when When the temperature in the reaction stage is higher than the temperature in the previous stage, it can be promoted. In addition, the vapor generated in the later stage can be used to heat the liquid in the previous stage. This can further improve the energy efficiency in the method of the present invention. Inventive method, sequentially introduce the liquid phase from the first section to the final row, and continuously introduce the catalyst and the aromatic hydroxy compound into the first section of the reactor (when it is in the liquid phase condition) And continuously introduce the aryl ester into at least one reaction section of the reactor (when it is in the gas phase or concentrated 312 / Invention _ Supplement) / 92-10 / 92120780 continuously around the ground with a dense phase, and Talent (carbon, depending on the alcohol). The use of steam and steam for the connection of several stirrers is economical, so the reaction in the later stage, the inverse rate in the zone. Time). Likewise, the aromatic hydroxy compound can be introduced into at least one reaction section of the reactor (when it is in a gas or liquid phase condition). In this way, diaryl carbonate or aromatic compounds are introduced into part or all of the separate reaction sections (when they are in the gas or liquid phase conditions), thereby promoting the evaporation of by-products (aliphatic alcohols) Effect, and the reaction equilibrium can be advantageously brought to the manufacturing system. Although not shown in the drawings, an internal coil can be provided in each reaction section, or an external jacket can be installed around the reaction zone, and a heat medium can be circulated in this coil and jacket to provide the energy required for the reaction In each reaction section of the reaction device of the present invention, a reboiler may be installed outside the reactor to provide heat to the reactor. Modifications of this effect are generally preferred in the present invention. The number of reaction zones (reaction sections) may be 2 or more, and is not particularly defined. However, even if the number is higher than necessary, the effect will gradually decrease. Therefore, the number may be usually 2 to 30, preferably 2 to 15. An external stirrer may be used to forcefully stir each of the reactors constituting the individual reaction zone therein and the individual reaction zone in the reactor, but this stirring is not necessary. Natural or convective flow or bubbling (which can occur by evaporation) may be sufficient to mix the reacting substances. External stirring systems that can be used here include, for example, those with stirring blades, those with pumping cycles, and those with steam or steam blowers. In the method for producing an aromatic carbonate according to the present invention, a carbonic acid or an aromatic hydroxy compound in a gas phase or a liquid phase is introduced into a partially or completely separated reaction zone in a reactor, and at the same time, a by-product (fat) is promoted. Group alcohols), thereby advantageously guiding the reaction equilibrium into the manufacturing system. In addition, most of the aliphatic alcohols formed in the later stage are taken out of the system and are not mixed with the reaction liquid in the previous stage. Therefore, the aliphatic alcohols will not accumulate in the earlier stage of 18 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 in the reaction zone (will not hinder the reaction). The method of the present invention generally does not require a solvent, but a solvent inert to the reaction (for example, selected from ethers, aliphatic hydrocarbons and aromatic hydrocarbons) may be used. The liquid phase taken out of the final reaction zone (reaction section) in the reaction device (reactor) is subjected to the purification component (for example, by distillation), and the desired alkylaryl carbonate and / or diaryl carbonate can be obtained. . Although it depends on the type and composition of the starting materials used, the reaction temperature in the method of the present invention can usually fall between 50 and 300 ° C, preferably between 100 and 250 ° C. At higher temperatures, the reaction rate will be higher, but by-products (alkyl aromatic ethers, etc.) will increase at high temperatures. Therefore, too high a temperature is disadvantageous for the method of the present invention. Depending on the type of starting material and composition in the reactor, the pressure inside the reactor can be changed. Generally speaking, it can be under pressure or reduced pressure (between 10 and 3000kPa, especially between 50 and 2000kPa (between 0. 5 and 20 atmospheres)). Generally, the catalyst is dissolved or dispersed in the starting material and supplied to the reaction zone (reactor). Based on the starting material to be supplied, the catalyst content is usually 0. 0 0 0 1 to 10 mole%, preferably 0. 0 0 1 to 5 mole%. If the catalyst content is too low, the reaction rate will be insufficient; but if it is too low, the content of by-products (alkyl aromatic ethers) may be increased. Depending on the other reaction conditions used, the average residence time of the liquid in the reaction device (reactor) can be changed, but usually falls below 0. Between 1 and 20 hours, preferably between 0. Between 3 and 10 hours. <Embodiments> Specific examples of the present invention will be further explained with reference to the following examples. However, 19 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 However, the present invention is limited to these without departing from the spirit thereof. Specific examples. [Example 1] This example uses two reactors, in which the reaction temperature in the later stage is the reaction temperature in the previous stage, and the reaction liquid system in the previous stage is heated by high-temperature steam generated for efficient energy utilization. The reaction system shown in Figure 1 with two 300 ml reactors was transesterified with phenol. The liquid level in each reactor was about 50%. Through the liquid feed line L 1, 94 g / h (1 mole / small phenol and 0. A catalyst of 6 g / hour (dibutyltin oxide) was supplied, and 90 g / hour (1 mole / hour) of dimethyl carbonate was supplied through the liquid line L1. The starting material temperature is 60 ° C. The pressure in the two reactors was 500 and the temperature in reactor 1 was 180 ° C. The temperature in reactor 2 was 0 ° C. The reaction liquid system is heated by the vapor generated in the reactor 2 in the heat exchanger 1, and then is sequentially guided to the line L3, the reactor 1, the line L4, and the device 2, and finally taken out through the liquid discharge line L5. The vapor generated in the reactor 1 is taken out through the vapor line V1. The steam generated in the device 2 is guided to the heat exchanger 1 via the line V 2 and is cooled therein; I is used to heat the starting material liquid, and then discharged through the line L 6 under this condition, so that the device is continuously driven 8 hours, and the reaction liquid discharged through L 5 was collected and analyzed for its composition. In this analysis, 5 can be detected. 9% by weight methylphenyl carbonate and 1. 6% by weight of diphenyl carbonate. The flow rate of line L 5 is 124 g / h, and the product methylphenyl carbonate 312 / Invention Specification (Supplement) / 92] 0/92120780. Higher than • Section production (when maintained)) Feed supply kPa, 200 medium plus reaction, discharge tube reaction " capacity, 〇 Pipeline passed and carbon 20 200408622 acid diphenyl ester equivalent to a total of 0. 0 5 7 mol / h. The heat required for the reaction is 4 9. 6 kilojoules / hour. The energy efficiency per mole is Π. 5χ1 (Γ4 Mol / kilojoule. [Example 2] This example uses three reactors, where the starting material liquid system in each stage generates high-temperature steam heating for efficient energy utilization in each stage. The reaction system with three 200 ml reactors shown in Figure 2 was transesterified with phenol. The liquid level in each reactor was maintained at about 50 ° / 〇. Via liquid feed Line L 1 was continuously supplied with the same amount of phenol and catalyst (dibutyltin oxide) as in Example 1, and was supplied through liquid feed line L 2 and separately through lines L 2 a, L 2 b, and L 2 c. The same amount of dimethyl carbonate as in Example 1. The flow rates through the lines L 2 a, L 2 b, and L 2 c were controlled to a ratio of 8/1 / 1. Regarding the conditions, all three reactors were set at 2 0 0 ° C and at 500 k Pa. The reaction liquid was sequentially guided to reactor 1, line L3, reactor 2, line L4, and reactor 3, and finally discharged through the liquid discharge line L5. Mix phenol and catalyst from line L1 with dimethyl carbonate from line L 2 a and then direct to heat exchanger 1 In the heat exchanger 1, the starting material liquid is heated by the vapor generated in the reactor 1, and then guided to the reactor 1. After being used for heat exchange, the vapor is condensed into a liquid at 80 ° C, and passed through This is discharged from line L 6. The steam generated at 200 ° C in reactor 2 is directed to heat exchanger 2 via line V 2 and this heated dimethyl carbonate from line L 2 b is followed by 21 312 / Invention Specification (Supplement) / 92 · 10 / 9212〇780 200408622 Cool to 16 6 ° C. The generated condensate is separated in a vapor-liquid separator 1 and returned to reactor 1 via line L7, while The uncondensed vapor was discharged through line V4. The heated dimethyl carbonate from line L2b was heated in heat exchanger 2 and the steam at 158 ° C was supplied to reactor 2. Via line V3, the The 200 ° C vapor generated in the reactor 3 is directed to the heat exchanger 3 (similar to the vapor of the reactor), and this heated dimethyl carbonate from the line L 2 c is then cooled to 150 ° C. The resulting condensate is separated in the vapor-liquid separator 2 and returned to the reaction via line L 8 2. At the same time, the uncondensed vapor is discharged through the line V 5. The heated dimethyl carbonate from the line L 2 c is heated in the heat exchanger 3, and the steam of 158 ° C is supplied to the reactor 3. Under this condition, the device was continuously driven for 8 hours, and the reaction liquid discharged through the line L 5 was collected and analyzed for composition. In this analysis, 6. 4 wt% fluorenyl carbonate and 1. 8% by weight of diphenyl carbonate. The flow rate through line L 5 was 127 g / h, and the products methylphenyl carbonate and diphenyl carbonate were equivalent to a total of 0. 0 6 5 mol / h. The heat required for the reaction is 5 3. 3 kilojoules / hour. The energy efficiency per mole is 12. 2xl (T4 Molars / kilojoules, and it is higher than that in Example 1. [Comparative Example 1] In this example, the reaction liquid system was operated in the same manner as in Example 2, but no self-generated steam was performed. Heat recovery. In the reaction system with three 200 ml reactors shown in Figure 3, 22 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 Transesterification of dimethyl carbonate with phenol Each reactor is about 50%. The same amount of phenol and catalyst (dibutyltin oxide) are continuously supplied through the liquid feed line L1 in FIG. 3, and via the line L 2 and separately through the line L 2 a, L 2 b and L 2 c, for the same amount of dimethyl carbonate. It will be controlled to the ratio of 5/2/3 through the lines L 2 a, L 2 b. Regarding the conditions, all three reactors are connected at 2 0 ° C, the reaction liquid is sequentially led to reactor 1, line L3, line L4 and reactor 3, and finally the vapor discharge line V2 produced in reactor 1 is discharged from the vapor discharge line V1 through the liquid discharge line , The steam generated in the reactor 2 is guided to be cooled to 160 ° C, and the generated condensate is returned to the line through the line L6, while the uncondensed The condensed vapor passes through the vapor discharge line line V 3, and the vapor generated in the reactor 3 is also cooled in the condenser as described above and cooled to 160 ° C, and is returned to the reactor 2 through the line L 7 liquid, while the uncondensed The condensed vapor was discharged through line V5. Under this condition, the device was continuously driven for 8 hours and collected; the discharged reaction liquid and its composition were analyzed. In this analysis, the weight% methylphenyl carbonate and 1. The flow rate of 8% by weight of diphenyl carbonate L 5 was 135 g / hour, and the product methyl J phenyl carbonate was equivalent to a total of 0. 0 7 2 moles / hour. Used to reverse 9 3. 4 kilojoules / hour. 312 / Instruction of the Invention (Supplement) / 92-10 / 92120780 Liquid level maintenance r is given to Example 1 from the liquid feed tube: given to the flow rate of Example 1 and L2c and at 500kPao reactor 2, tube L 5 discharge. Exhausted. Passed through the condenser and cooled back to reactor 1 V 4 and discharged. Guided by means of which the generated condensate can be detected by the steam exhaust pipe through line L5 6. 7 esters. The heat required for good esters and carbonates through the pipeline 23 200408622 The energy efficiency per mole product is 7. 7xl (T4 Moore / kilojoule, and it is confirmed that this is lower than in the previous example. [Comparative Example 2] This example uses three reactors connected in series, in which the starting materials phenol and dimethyl carbonate are countercurrent to each other In a reaction system with three 200-ml pressure cookers connected in series as shown in Fig. 4, dimethyl carbonate was transesterified with phenol in a countercurrent contact manner with each other. The liquid level in each reactor was maintained at About 50%. The same amount of phenol and catalyst (dibutyltin oxide) as in Example 1 are continuously supplied through the liquid feed line L1 in FIG. 4, and the same as in Example 1 are supplied through the liquid feed line L 2. The same amount of dimethyl carbonate, where this is completely gasified, and then supplied to reactor 3 via line V 1 a. The pressure in the reactor is controlled so that it can be increased from reactor 1 to reactor 3 for steam Flow. Specifically, the pressure in reactor 1 is 500 kPa, the pressure in reactor 2 is 550kPa, and the pressure in reactor 3 is 600kPa. Regarding the reaction temperature, reactor 3 is heated The temperature can reach 200 ° C, but reactor 1 and the reaction The reactor 2 is not heated. The reaction liquid is sequentially passed through the reactor 1, the line L 3 a, the reactor 2, the line L 3 b, and the reactor 3, and finally discharged through the liquid discharge line L 3 c. Through the line V 1 b, The 200 ° C steam generated in the reactor 3 is supplied to the reactor 2; the steam generated in the reactor 2 is supplied to the reactor 1 through the line V 1 c; and the reactor 1 is supplied through the line V 2 The generated vapor is discharged. Under this condition, the device is continuously driven for 8 hours, and the reaction liquid discharged through the line L 3 c is collected and its composition is analyzed. In this analysis, 7 can be detected. 4 24 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 wt% methylphenyl carbonate and 1. 6% by weight of diphenyl carbonate. The flow rate through line L 3 c was 149 g / h, and the products methylphenyl carbonate and diphenyl carbonate corresponded to a total of 0. 0 8 4 mol / h. The heat required for the reaction is 7 3. 2 kilojoules / hour. The energy efficiency per mole is 11. 5x10 _4 moles / kilojoule, and it was confirmed that it was similar to that in Example 1, but lower than that in Example 2. As in Examples 1 and 2 and Comparative Examples 1 and 2, according to the method of the present invention, the liquid phase part passes through a plurality of separate reaction zones in series, and the light ends are distilled off, so the light ends (for example, The initial alcohol (equivalent aliphatic alcohol) can be minimized without any additional and significant equipment costs, and in addition, a sufficient reaction volume can be easily ensured. Therefore, the method according to the present invention makes it possible to continuously produce an aromatic polycarbonate, and its productivity and selectivity are equally high. In addition, in the method of the present invention, the reaction temperature, reaction pressure, residence time, and other conditions can be easily selected over a wide range, and yet another advantage of the present invention is that extremely simplified equipment can be used to achieve extremely high productivity and Product of selectivity. [Example 3] Using a 600 ml reactor divided into three reaction sections (as shown in the structure of Fig. 5), according to the reaction system shown in Fig. 6, dimethyl carbonate was transesterified with phenol. As shown in Figure 6, through the liquid feed line L1, 94 g / h (1 mol / h) of phenol and 0 were continuously supplied. 6 g / h catalyst (dibutyltin oxide), and supplied 90 g / h (1 mol / h) via liquid feed line L 2 and separately via lines L 2 a, L 2 b, and L 2 c The dihydrazone carbonate is intended for each 25 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 reaction section. Control the flow rate through lines L 2 a, L 2 b and L 2 c to 5/2/3 ° The pressure in the reactor is 500 kPa, and the temperature is 200 ° C The reaction liquid overflows from the first section To the second section and to the third section, they are discharged through the liquid discharge line L 3. It is discharged through the steam generated in the first section of the steam discharge line V1. Via line V2, the steam generated in the second zone is guided to the condenser and cooled to 160 ° C, and the raw condensate is returned to the first zone via line L4, while the uncondensed steam is discharged from the steam discharge line V 4 is discharged. Via line V 3, the steam generated in the third zone is also guided to the condenser and cooled to 160 ° C, and the raw condensate is returned to the second zone through line L 5 while the uncondensed steam is removed from the steam. The discharge line V 5 is discharged. Under this condition, the apparatus was continuously driven for 8 hours, and the reaction liquid discharged through the tube was collected and analyzed for its composition. In this analysis, it is possible to detect the weight% fluorenylphenyl carbonate and 1. 8% by weight of diphenyl carbonate. The flow rate through L 3 was 135 g / hour, and the products methylphenyl carbonate and carbyl phenyl ester were equivalent to a total of 0. 0 7 2 moles / hour. Required for reaction is 9 3. 4 kilojoules / hour. The liquid level in the reactor was maintained at an energy efficiency of about 7 per mole of product. 7χ1 (Γ4 mol / kilojoule. [Comparative Example 3] Continuous supply of 94 g / hr (1 mol / hr) phenol, 0.6 g / λ catalyst (dibutyltin oxide), and 90 g / hr (1 mole / hour) carbonate into a 500 ml pressure cooker, while the liquid level in the reactor was about 60%. The pressure in the reactor was 500 kPa, and the temperature was 312. / Invention specification (Supplement) / 92_ 10 / 9212〇780 ratio 0 make and connect, the gas meridian line L3 of the middle part of the gas flow through the segment 16. 7 Pipeline acid heat 50 0% ° When the temperature is maintained at 200 26 200408622 ° c. The steam generated in the reactor is continuously discharged. The reaction was continued for 8 hours, and the reaction liquid was collected and analyzed for its composition. In this analysis, 5. 9% by weight of methylphenyl carbonate and 1.6% by weight of diphenyl carbonate. The discharge rate of the reaction liquid was 12 2 g / hour. The products methylphenyl carbonate and diphenyl carbonate are equivalent to a total of 0. 0 5 7 mol / h. The heat required for the reaction is 8 2. 0 kilojoules / hour. The energy efficiency per mole is 7. Οχ 1 (Γ4 Mol / kilojoule, and the productivity per energy is smaller than that in Example 3. [Comparative Example 4] The reactor of Fig. 7 was used in which the liquid phase was divided into three sections, but there was no higher partition wall (The vapor phase is divided by this.) Therefore, the vapor phase forms a continuous vapor phase, and dimethyl carbonate is transesterified with phenol. Specifically, the same amount of phenol and catalyst as in Example 3 are supplied through the liquid feed line L1 ( Dibutyltin oxide), and through the liquid feed lines L 2 a, L 2 b, and L 2 c, the same amount of dimethyl carbonate as in Example 3 is supplied to the individual reaction section. However, it will be passed through the line L 2 a, The flow rates of L 2 b and L 2 c are controlled to a ratio of 6/2/2. The pressure in the reactor is 500 kPa and the temperature is 200 ° C. The reaction liquid overflows from the first section to the second section. And to the third section, and then discharged through the liquid discharge line L 3. The vapor generated in each reaction section becomes uniform in the reactor and is discharged through the vapor discharge line V1. Under this condition, the device is continuously Drive for 8 hours and collect the reaction liquid discharged through line L3 and analyze its group . This analysis can detect up to 6. 5 wt% methylphenyl carbonate and 1. 8% by weight of diphenyl carbonate. The flow rate through line 27 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 L 3 is 12 2 g / hour, and the products fluorenyl carbonate and diphenyl carbonate are equivalent to a total of 0. 0 6 2 moles / hour. The heat required for the reaction is 8 1. 9 kilojoules / hour. The energy efficiency per mole is 7. 6xl0_4 Moore / kilojoule. The productivity per energy is larger than that in Comparative Example 1, but still smaller than that in Example 3. [Example 4] Using the reactor of Fig. 5, according to the reaction system shown in Fig. 8, dimethyl carbonate was transesterified with phenol. As shown in FIG. 8, the same amount of phenol and catalyst (dibutyltin oxide) as in Example 3 were continuously supplied through the liquid feed line L1, and through the liquid feed line L 2 and separately through the lines L 2 a, L 2 b and L 2 c, and dimethyl carbonate was supplied. The flow rates through the lines L 2 a, L 2 b and L 2 c are controlled to a ratio of 8/1 /: 1. The phenol and the catalyst from the line L 1 are mixed with the dimethyl carbonate from the line L 2 a and supplied to the heat exchanger E 1 through the line L 3. The heating is started by the steam generated in the first section of the reactor. Material, and then enter the first section of the reactor. After undergoing heat exchange, the vapor became condensate at 80 ° C and was discharged through line L6. The pressure in the reactor was 500 kPa and the temperature was 200 ° C. The reaction liquid is allowed to overflow from the first section to the second section and to the third section, and is then discharged through the liquid discharge line L5. The 200 ° C vapor generated in the second section was introduced into the heat exchanger E 2 via the line V 2, and then the dimethyl carbonate from the line L 2 b was heated and cooled to 16 6 ° C. The resulting condensate is separated in a vapor-liquid separator and returned to the first section via line L4a, while uncondensed vapor is passed through line 28 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622 line V 4 is discharged. Dimethyl carbonate from line L 2 b is heated in heat exchanger E 2 and steam at 158 ° C is supplied to the middle of the reactor. 0 2 generated in the third section is passed through line V3. 0 ° C steam is introduced into the exchanger E 3 (similar to the steam in the second section), and this heated dimethyl carbonate from the tube is then cooled to 150 ° C. The generated condensate is separated in a vapor-liquid and returned to the line via line L 4b, while uncondensed vapor is discharged through line V5. The heated dimethyl carbonate S from the line L 2 c is heated in the heat E 3 and the steam is supplied to the third section. Under this condition, the apparatus was continuously driven for 8 hours, and the reaction liquid discharged through it was collected and analyzed for its composition. In this analysis, it is possible to detect the weight% fluorenyl carbonate and 1. 8% by weight of diphenyl carbonate. The flow rate of L 5 was 127 g / h, and the products methylphenyl carbonate and phenyl carbonate were equivalent to a total of 0. 0 6 5 mol / h. Required for reaction is 5 3. 3 kilojoules / hour. The energy efficiency per mole is 12. 2 Οχ 1 (Γ4 Mol / kJ is high. [Comparative Example 5] Three 200 ml pressure cookers were connected in series, and according to the system of Fig. 9, the difluorene carbonate was transesterified with phenol in countercurrent contact with each other The starting material phenol and dimethyl carbonate are in countercurrent contact with each other. As shown in FIG. 9, the same amount of phenol and catalyst (dibutyltin oxide) as the actual amount are continuously supplied to the reactor R 1 through the liquid feed line L1. Heated by the second section heated by 312 / Instruction Manual (Supplement) / 92-10 / 92120780. Line L 2 c Body separation Second zone t converter 1 58 ° C Lines L5 to 6. 4 Pass the heat of dicarbonate through the pipeline and react with the ester. Its example 1 is from liquid 29 200408622 feed line L 2 and supplies dimethyl carbonate to heat exchanger E 1 where it is gasified and then supplied to reactor R 3 via line V 1 a. The pressure in the reactor is controlled so that it can be raised from reactor R 1 towards reaction i for the flow of steam. Specifically, the pressure in the reactor R1 is kPa, the pressure in the reactor R2 is 550 kPa, and the force in the reactor R3 is 600 kPa. Regarding the reaction temperature, the reactor R 3 was heated so that the degree could be 200 ° C, but the reactors R1 and R2 were not heated. The reaction liquid was sequentially passed through the reactor R1, the line L3a, the reactor line L3b, and the reactor R3, and finally through the liquid discharge line L3c to the line V1b, and the reactor R3 The generated 200 ° C steam is supplied to the reactor 2; the steam generated in the reactor R 2 is passed to the reactor R1 through a line V1 c; and the steam generated in the reactor R1 is discharged through a line V2. Under this condition, the apparatus was continuously driven for 8 hours, and the reaction liquid discharged through the tube green was collected and analyzed for its composition. In this analysis, weight percent methylphenyl carbonate and 1. 6% by weight of diphenyl carbonate. The flow rate through L 3 c is 149 g / h, and the products methylphenyl carbonate and carbyl phenyl ester are equivalent to a total of 0. 0 8 4 mol / h. Required for reaction is 7 3. 2 kilojoules / hour. The energy efficiency per mole is 11. 47xl (T4 mole / kilojoule is high, but lower than in Example 4. As in Examples 3 and 4 and Comparative Examples 3, 4 and 5, the special reactor of the present invention allows the permeation of aromatic hydroxy compounds and carbonic acid The transesterification of dialkyl esters makes it possible to continuously produce aromatic carbonates, and its energy efficiency is high 312 / Invention Specification (Supplements) / 92-10 / 92120780 Completely R3 500 pressure and temperature R2. To supply steam. L3c 7. 4 Pipeline acid heat, and special reaction 30 200408622 In general, the transesterification reaction is an equilibrium reaction, which is greatly transferred to the starting system, and its reaction rate is low. The present invention has been described in detail with reference to its specific examples, but those skilled in the art will understand that many changes and modifications can be made without departing from the spirit and scope thereof. This application is filed on August 12, 2002. This patent application is filed (Application No. 2 0 2-2 3 4 8 8 4) and August 13, 2002 Based on Japanese Patent Application (Application No. 2002-2 3 5 3 8 5), and its contents are incorporated in this case for reference. <Industrial Applicability> According to the method of the present invention, the liquid phase part passes through a plurality of separate reaction zones in series, and the light ends are distilled off, so the light ends (e.g., aliphatic which may cause reaction equilibrium to the initial system) Alcohol) can be minimized without any additional and significant equipment costs, and in addition, a sufficient reaction volume can be easily ensured. Therefore, the method according to the present invention makes it possible to continuously produce an aromatic polycarbonate, and its productivity and selectivity are likewise high. In addition, in the method of the present invention, the reaction temperature, reaction pressure, residence time, and other conditions can be easily selected over a wide range, and yet another advantage of the present invention is that extremely simplified equipment can be used to achieve extremely high productivity and Product of selectivity. [Brief Description of the Drawings] Fig. 1 is a schematic sectional view showing an embodiment of a reaction device having two reactors connected in series. Fig. 2 is a schematic cross-sectional view showing an example of a reaction device with three reactors connected in series 31 312 / Invention Specification (Supplement) / 92-10 / 92120780 200408622. Fig. 3 is a schematic sectional view showing a reactor structure and a reaction system of Comparative Example 1. Fig. 4 is a schematic sectional view showing a reactor structure and a reaction system of Comparative Example 2. Fig. 5 is a schematic sectional view showing an embodiment of a reactor structure divided into three reaction zones inside. Fig. 6 is a schematic sectional view showing a reactor structure and a reaction system of Example 3. Fig. 7 is a schematic sectional view showing a reactor structure and a reaction system of Comparative Example 4. Fig. 8 is a schematic sectional view showing the structure of a reactor and a reaction system of Example 4. Fig. 9 is a schematic sectional view showing a reactor structure and a reaction system of Comparative Example 5. (Description of component symbols) 1, 2, 3, R 1, R 2, R 3 Reactor (reaction tank) L 1, L 2 Starting material feed line L 3, L 4, L 5, L 6, L 7 , L 8 liquid discharge line V 1, V 2, V 3, V 4, V 5 vapor discharge line E 1, E 2, E 3 heat exchanger A 1, A 2 lower compartment
Bl、B2 較高隔間 32 312/發明說明書(補件)/92-10/92120780Bl, B2 Higher compartment 32 312 / Invention specification (Supplement) / 92-10 / 92120780