JP4386609B2 - Continuous production method of polycarbonate - Google Patents

Description

【００１１】
【数６】

【００１２】
【数７】

【００１３】
【数８】

【００１４】
（但し、式中、OH₀はプレポリマーの目標水酸基末端比率（モル％）、Mn₀はプレポリマーの目標数平均分子量、OH_aはプレポリマーの実測水酸基末端比率、Mn_aはプレポリマーの実測数平均分子量、Mn_bは添加するジアリールカーボネートの分子量、Mn_cは添加する芳香族ジヒドロキシ化合物の分子量、X_biは時点ｉで添加するジアリールカーボネートのモル分率量、X_ciは時点ｉで添加する芳香族ジヒドロキシ化合物のモル分率、W_aiは時点ｉで反応器を通過するプレポリマーの重量（ｇ/hr）、Wbiは時点ｉで添加するジアリールカーボネートの重量（ｇ/hr）、Wciは時点ｉで添加する芳香族ジヒドロキシ化合物の重量（ｇ/hr）である。aは試行錯誤で求められる装置係数であり、0.5〜1.0の範囲にある。）
を提供するものである。
【００１５】
以下に本発明について詳細に説明する。
本発明において、芳香族ジヒドロキシ化合物とは、ＨＯ−Ａｒ−ＯＨで示される化合物である（式中、Ａｒは２価の芳香族基を表す。）。
芳香族基Ａｒは、例えば好ましくは−Ａｒ¹−Ｙ−Ａｒ²−で示される２価の芳香族基である（式中、Ａｒ¹及びＡｒ²は、各々独立にそれぞれ炭素数５〜７０を有する２価の炭素環式又は複素環式芳香族基を表し、Ｙは炭素数１〜３０を有する２価のアルカン基を表す。）。
【００１６】
２価の芳香族基Ａｒ¹、Ａｒ²において、１つ以上の水素原子が、反応に悪影響を及ぼさない他の置換基、例えば、ハロゲン原子、炭素数１〜１０のアルキル基、炭素数１〜１０のアルコキシ基、フェニル基、フェノキシ基、ビニル基、シアノ基、エステル基、アミド基、ニトロ基などによって置換されたものであっても良い。
複素環式芳香族基の好ましい具体例としては、１ないし複数の環形成窒素原子、酸素原子又は硫黄原子を有する芳香族基を挙げる事ができる。
２価の芳香族基Ａｒ¹、Ａｒ²は、例えば、置換又は非置換のフェニレン、置換又は非置換のビフェニレン、置換または非置換のピリジレンなどの基を表す。ここでの置換基は前述のとおりである。
２価のアルカン基Ｙは、例えば、下記化１で示される有機基である。
【００１７】
【化１】

【００１８】
（式中、Ｒ¹、Ｒ²、Ｒ³、Ｒ⁴は、各々独立に水素、炭素数１〜１０のアルキル基、炭素数１〜１０のアルコキシ基、環構成炭素数５〜１０のシクロアルキル基、環構成炭素数５〜１０の炭素環式芳香族基、炭素数６〜１０の炭素環式アラルキル基を表す。ｋは３〜１１の整数を表し、Ｒ⁵およびＲ⁶は、各Ｘについて個々に選択され、お互いに独立に、水素または炭素数１〜６のアルキル基を表し、Ｘは炭素を表す。また、Ｒ¹、Ｒ²、Ｒ³、Ｒ⁴、Ｒ⁵、Ｒ⁶において、一つ以上の水素原子が反応に悪影響を及ぼさない範囲で他の置換基、例えばハロゲン原子、炭素数１〜１０のアルキル基、炭素数１〜１０のアルコキシ基、フェニル基、フェノキシ基、ビニル基、シアノ基、エステル基、アミド基、ニトロ基等によって置換されたものであっても良い。）
このような２価の芳香族基Ａｒとしては、例えば、下記化２で示されるものが挙げられる。
【００１９】
【化２】

【００２０】
（式中、Ｒ⁷、Ｒ⁸は、各々独立に水素原子、ハロゲン原子、炭素数１〜１０のアルキル基、炭素数１〜１０のアルコキシ基、環構成炭素数５〜１０のシクロアルキル基またはフェニル基であって、ｍおよびｎは１〜４の整数で、ｍが２〜４の場合には各Ｒ⁷はそれぞれ同一でも異なるものであってもよいし、ｎが２〜４の場合には各Ｒ⁸はそれぞれ同一でも異なるものであってもよい。）
さらに、２価の芳香族基Ａｒは、−Ａｒ¹−Ｚ−Ａｒ²−で示されるものであっても良い。
（式中、Ａｒ¹、Ａｒ²は前述の通りで、Ｚは単結合又は−Ｏ−、−ＣＯ−、−Ｓ−、−ＳＯ₂−、−ＳＯ−、−ＣＯＯ−、−ＣＯＮ（Ｒ¹）−などの２価の基を表す。ただし、Ｒ¹は前述のとおりである。）
このような２価の芳香族基Ａｒとしては、例えば、下記化３で示されるものが挙げられる。
【００２１】
【化３】

【００２２】
（式中、Ｒ⁷、Ｒ⁸、ｍおよびｎは、前述のとおりである。）
さらに、２価の芳香族基Ａｒの具体例としては、置換または非置換のフェニレン、置換または非置換のナフチレン、置換または非置換のピリジレン等が挙げられる。ここでの置換基は前述のとおりである。
本発明で用いられる芳香族ジヒドロキシ化合物は、単一種類でも２種類以上でもかまわない。芳香族ジヒドロキシ化合物の代表的な例としてはビスフェノールＡが挙げられる。
本発明で用いられるジアリールカーボネートは、下記化４で表される。
【００２３】
【化４】

【００２４】
（式中、Ａｒ³、Ａｒ⁴はそれぞれ１価の芳香族基を表す。）
Ａｒ³及びＡｒ⁴は、１価の炭素環式又は複素環式芳香族基を表すが、このＡｒ³、Ａｒ⁴において、１つ以上の水素原子が、反応に悪影響を及ぼさない他の置換基、例えば、ハロゲン原子、炭素数１〜１０のアルキル基、炭素数１〜１０のアルコキシ基、フェニル基、フェノキシ基、ビニル基、シアノ基、エステル基、アミド基、ニトロ基などによって置換されたものであっても良い。Ａｒ³とＡｒ⁴は同じものであっても良いし、異なるものであっても良い。
１価の芳香族基Ａｒ³及びＡｒ⁴の代表例としては、フェニル基、ナフチル基、ビフェニル基、ピリジル基を挙げる事ができる。これらは、上述の１種以上の置換基で置換されたものでも良い。
好ましいＡｒ³及びＡｒ⁴としては、それぞれ例えば、下記化５などが挙げられる。
【００２５】
【化５】

【００２６】
ジアリールカーボネートの代表的な例としては、下記化６で示される置換または非置換のジフェニルカーボネート類を挙げる事ができる。
【００２７】
【化６】

【００２８】
（式中、Ｒ⁹及びＲ¹⁰は、各々独立に水素原子、炭素数１〜１０を有するアルキル基、炭素数１〜１０を有するアルコキシ基、環構成炭素数５〜１０のシクロアルキル基又はフェニル基を示し、ｐ及びｑは１〜５の整数で、ｐが２以上の場合には、各Ｒ⁹はそれぞれ異なるものであっても良いし、ｑが２以上の場合には、各Ｒ¹⁰は、それぞれ異なるものであっても良い。）
このジフェニルカーボネート類の中でも、非置換のジフェニルカーボネートや、ジトリルカーボネート、ジ−ｔ−ブチルフェニルカーボネートのような低級アルキル置換ジフェニルカーボネートなどの対称型ジアリールカーボネートが好ましいが、特にもっとも簡単な構造のジアリールカーボネートである非置換のジフェニルカーボネートが好適である。
【００２９】
これらのジアリールカーボネート類は単独で用いても良いし、２種以上を組み合わせて用いても良い。
芳香族ジヒドロキシ化合物とジアリールカーボネートとの使用割合（仕込比率）は、目標とするポリカーボネートの分子量や末端基比率、用いられる芳香族ジヒドロキシ化合物とジアリールカーボネートの種類、重合温度その他の重合条件によって異なるが、ジアリールカーボネートが芳香族ジヒドロキシ化合物１モルに対して、通常０．９〜２．５モル、好ましくは０．９５〜２．０モル、より好ましくは０．９８〜１．５モルの割合で用いられる。製造するポリカーボネートの目標分子量や水酸基末端比率によって、芳香族ジヒドロキシ化合物とジアリールカーボネートとの仕込み比率を変更する方法が一般的に用いられている。しかしながら、仕込み比率を変更した場合には、最終的に目標とする分子量や水酸基末端比率を有するポリカーボネートが得られるまでに、多くの切り替えロスを伴うことから好ましくない。本発明においては、製造するポリカーボネートの目標分子量や水酸基末端比率が変更になっても、両者の仕込み比率を変更せずに一定のままでポリカーボネートを製造することが好ましい。
【００３０】
溶融重縮合反応は、触媒を加えずに実施する事ができるが、重合速度を高めるため、必要に応じて触媒の存在下で行われる。重合触媒としては、この分野で用いられているものであれば特に制限はないが、水酸化リチウム、水酸化ナトリウム、水酸化カリウム、水酸化カルシウムなどのアルカリ金属またはアルカリ土類金属の水酸化物類；水素化アルミニウムリチウム、水素化ホウ素ナトリウム、水素化ホウ素テトラメチルアンモニウムなどのホウ素やアルミニウムの水素化物のアルカリ金属塩、アルカリ土類金属塩、第四級アンモニウム塩類；水素化リチウム、水素化ナトリウム、水素化カルシウムなどのアルカリ金属またはアルカリ土類金属の水素化合物類；リチウムメトキシド、ナトリウムエトキシド、カルシウムメトキシドなどのアルカリ金属またはアルカリ土類金属のアルコキシド類；リチウムフェノキシド、ナトリウムフェノキシド、マグネシウムフェノキシド、ＬｉＯ−Ａｒ−ＯＬｉ、ＮａＯ−Ａｒ−ＯＮａ（Ａｒはアリール基）などのアルカリ金属またはアルカリ土類金属のアリーロキシド類；酢酸リチウム、酢酸カルシウム、安息香酸ナトリウムなどのアルカリ金属またはアルカリ土類金属の有機酸塩類；酸化亜鉛、酢酸亜鉛、亜鉛フェノキシドなどの亜鉛化合物類；酸化ホウ素、ホウ酸、ホウ酸ナトリウム、ホウ酸トリメチル、ホウ酸トリブチル、ホウ酸トリフェニル、（Ｒ¹Ｒ²Ｒ³Ｒ⁴）ＮＢ（Ｒ¹Ｒ²Ｒ³Ｒ⁴）で表されるアンモニウムボレート類、（Ｒ¹Ｒ²Ｒ³Ｒ⁴）ＰＢ（Ｒ¹Ｒ²Ｒ³Ｒ⁴）で表されるホスホニウムボレート類（Ｒ¹、Ｒ²、Ｒ³、Ｒ⁴は前記化１の説明通りである。）などのホウ素の化合物類；酸化ケイ素、ケイ酸ナトリウム、テトラアルキルケイ素、テトラアリールケイ素、ジフェニル−エチル−エトキシケイ素などのケイ素の化合物類；酸化ゲルマニウム、四塩化ゲルマニウム、ゲルマニウムエトキシド、ゲルマニウムフェノキシドなどのゲルマニウムの化合物類；酸化スズ、ジアルキルスズオキシド、ジアルキルスズカルボキシレート、酢酸スズ、エチルスズトリブトキシドなどのアルコキシ基またはアリーロキシ基と結合したスズ化合物、有機スズ化合物などのスズの化合物類；酸化鉛、酢酸鉛、炭酸鉛、塩基性炭酸塩、鉛及び有機鉛のアルコキシドまたはアリーロキシドなどの鉛の化合物；第四級アンモニウム塩、第四級ホスホニウム塩、第四級アルソニウム塩などのオニウム化合物類；酸化アンチモン、酢酸アンチモンなどのアンチモンの化合物類；酢酸マンガン、炭酸マンガン、ホウ酸マンガンなどのマンガンの化合物類；酸化チタン、チタンのアルコキシドまたはアリーロキシドなどのチタンの化合物類；酢酸ジルコニウム、酸化ジルコニウム、ジルコニウムのアルコキシド又はアリーロキシド、ジルコニウムアセチルアセトンなどのジルコニウム化合物類などの触媒を挙げる事ができる。
【００３１】
触媒を用いる場合、これらの触媒は１種だけで用いても良いし、２種以上を組み合わせて用いても良い。また、これらの触媒の使用量は、原料の芳香族ジヒドロキシ化合物１００重量部に対して、通常１０^-8〜１重量部、好ましくは１０^-7〜１０^-1重量部の範囲で選ばれる。
本発明の方法は、ポリカーボネートプレポリマー製造工程Ａ、ポリカーボネートプレポリマーの水酸基末端比率制御工程Ｂ及び後重合工程Ｃからなる製造方法であり、一系列のＡ工程に対して複数のＢ工程及び／もしくは複数のＣ工程からなることを特徴としている。各工程の反応器としては、ポリカーボネートの反応器として広く使用されているものが使用でき、例えば、攪拌槽型反応器、薄膜反応器、遠心式薄膜蒸発反応器、表面更新型二軸混練反応器、二軸横型攪拌反応器、濡れ壁式反応器、自由落下させながら重合する多孔板型反応器、ワイヤーに沿わせて落下させながら重合するワイヤー付き多孔板型反応器等が用いられる。これらの反応器の材質は特に限定されないが、鉄を２０％以上含む材質が好ましく、特にＳＵＳ３０４，ＳＵＳ３１６，ＳＵＳ３１６Ｌが好ましく用いられる。
【００３２】
ポリカーボネートプレポリマー製造工程Ａは、芳香族ジヒドロキシ化合物とジアリールカーボネート等の原材料の計量・供給工程、これらを溶解する原料溶解槽及び反応器とからなり、原料溶解槽及びもしくは反応器の後にフィルターを設置しても良い。原料溶解槽はバッチ式で行ってもよく、連続で実施してもよいが、前者が好ましい。その際、複数の原料溶解槽を並列用いて切り替えながら使用することが好ましい。反応器は単独でも良く、複数の反応器を直列に連結してもよい。これら原料溶解槽及び反応器としては、前者の内、特に攪拌槽型反応器が好ましく用いられる。
【００３３】
原料の芳香族ジヒドロキシ化合物とジアリールカーボネートは、計量後、原料溶解混合槽に移送されるが、芳香族ジヒドロキシ化合物およびジアリールカーボネートは、予め窒素等の不活性ガスを用いて真空置換しておくことがこのましい。また、芳香族ジヒドロキシ化合物及びジアリールカーボネートのどちらか一方を、もしくは両方を酸素が除去された溶融状態で原料溶解混合槽に移送する方法も好ましい。特に芳香族ジヒドロキシ化合物もしくはジアリールカーボネートの製造工場が隣接する場合、配管等で接続して直接原料溶解混合槽に移送することが好ましい。原料溶解槽の温度は１１５〜２２０℃であるが、好ましくは１２０〜２００℃、さらに好ましくは１４０〜１９０℃の範囲にある。該原料溶解混合槽では、芳香族ジヒドロキシ化合物とジアリールカーボネートとを単に溶解混合するだけでなく、エステル交換反応を進行させてもよい。反応率としては５〜９５％の範囲で調整し、好ましくは３０〜９２％、特に好ましくは５０〜９０％の範囲で調整した後、後工程、例えば貯槽や反応器に移送することが好ましい。
【００３４】
ポリカーボネートプレポリマー製造工程Ａでの反応温度は、通常１００〜３５０℃、好ましくは１５０〜２９０℃の温度の範囲で選ばれる。好ましくは１８０〜２７０℃、特に好ましくは２００〜２５０℃の範囲にある。得られるポリカーボネートプレポリマーの数平均分子量は１５００〜７５００の範囲にある。好ましい数平均分子量の範囲としては１８００〜６０００、より好ましくは２０００〜５０００の範囲にある。反応の進行にともなって、芳香族モノヒドロキシ化合物が生成してくるが、これを反応系外へ除去する事によって反応速度が高められる。従って、窒素、アルゴン、ヘリウム、二酸化炭素や低級炭化水素ガスなど反応に悪影響を及ぼさない不活性なガスを導入して、生成してくる該芳香族モノヒドロキシ化合物をこれらのガスに同伴させて除去する方法や、減圧下に反応を行う方法などが好ましく用いられる。好ましい反応圧力は、重合初期には１０ｍｍＨｇ〜常圧の範囲にある。
【００３５】
ポリカーボネートプレポリマー製造工程Ａで製造されたプレポリマーは、ポリカーボネートプレポリマーの水酸基末端比率制御工程Ｂへ送られる。工程Ｂは、ポリカーボネートプレポリマーの分子量と水酸基末端比率を測定し、芳香族ジヒドロキシ化合物もしくはジアリールカーボネートを添加して反応させることで、該水酸基末端比率を制御する工程である。 工程Ｂは一系列でも良いし、複数の系列であっても良い。分子量と水酸基末端比率との関係が全く異なる複数の水酸基末端比率を有するポリカーボネートを同時に連続製造する場合には、対応した複数の系列が必要である。工程Ｂでの反応器は一系列当たり１〜２基であり、添加される芳香族ジヒドロキシ化合物もしくはジアリールカーボネートとの反応を充分行わせるために、攪拌槽型反応器が好ましく用いられる。反応温度は、通常１００〜３５０℃、好ましくは１５０〜２９０℃の温度の範囲で選ばれる。好ましくは１８０〜２７０℃、特に好ましくは２００〜２５０℃の範囲にある。好ましい反応圧力は１０ｍｍＨｇ〜常圧の範囲にある。
【００３６】
本発明において、ポリカーボネートプレポリマーの水酸基末端比率は、一般に５〜９５モル％の範囲にあり、この範囲の目標水酸基末端比率に制御される。ポリマーアロイ等の目的で水酸基末端比率が５０％以上のポリカーボネートを製造する場合には、プレポリマー段階での目標水酸基末端比率が５０〜９５モル％にある。光学用途や一般射出成形、押出成形用に適した水酸基末端比率が５０％以下のポリカーボネートを製造する場合には、プレポリマー段階での目標水酸基末端比率が５〜５０モル％にある。後者の場合、ポリカーボネートプレポリマーの水酸基末端比率を５〜５０モル％の範囲に設定することが好ましい。特に２０〜４８モル％が好ましい。
【００３７】
反応器から抜き出されたポリカーボネートプレポリマーの分子量は、例えば、溶液粘度、GPC、溶融粘度等によって測定することができる。一般に溶液粘度法は、簡便で迅速に分子量の測定ができるが、得られる分子量が粘度平均分子量であるので、予め数平均分子量と粘度平均分子量との相関を求めておくことが必要である。また、溶融粘度法では、例えば配管やサンプリング口に粘度計を設置することで連続的且つ自動的に測定できる。その際、溶融粘度と数平均分子量の相関を求めておくことが必要である。測定頻度は、高くするほど得られるポリカーボネートの分子量変動が小さくなり好ましいが、分子量の測定方法によっても制約を受ける。一般的には、数秒〜６時間の間隔にある。
【００３８】
反応器から抜き出されたポリカーボネートプレポリマーの水酸基末端比率は、例えば、ＮＭＲ、IR、比色（チタン法）、液体クロマトグラフィー等によって測定できる。これらの中でＮＭＲ法及び液体クロマトグラフィーが迅速であり精度も高く好ましい。測定頻度は、高くするほど得られるポリカーボネートの水酸基末端比率変動が小さくなり好ましいが、水酸基末端比率の測定方法によっても制約を受ける。一般的には、数秒〜６時間の間隔にある。本発明の製造方法において、ポリカーボネートの末端は、出発原料の芳香族ジヒドロキシ化合物に由来する水酸基末端（Ａ）、出発原料のジアリールカーボネートに由来するアリール基末端（Ｂ）がある。水酸基末端比率（モル％）は、下記式５で求められる。
【００３９】
【数９】

【００４０】
例えば、出発原料としてビスフェノール−Ａ及びジフェニルカーボネートを用いた場合には、ビスフェノール−Ａの水酸基末端（Ａ）とフェニル末端（Ｂ）のモル数とから水酸基末端比率が求められる。
本発明においては、ポリカーボネートプレポリマーの目標水酸基末端比率に制御するために、該プレポリマーにジアリールカーボネートもしくは芳香族ジヒドロキシ化合物を添加して反応させる。プレポリマーの水酸基末端比率を制御するために添加される芳香族ジヒドロキシ化合物やジアリールカーボネートとしては、出発原料として用いた芳香族ジヒドロキシ化合物もしくはジアリールカーボネートと同一の物を用いることが好ましい。しかし、他の芳香族ジヒドロキシ化合物やジアリールカーボネートや、両末端にアリールカーボネート末端を有するポリカーボネートポリマーやオリゴマー、両末端に水酸基末端を有するポリカーボネートポリマーやオリゴマーを使用することもできる。また、これらの添加は連続的に行われるため、添加方法としてはポンプを用いることが好ましく、特にプレポリマーの液中に供給することが好ましい。これら化合物は固体状態でも添加することができるが、ポンプを用いて溶融状態もしくは反応に影響を及ぼさない溶剤に溶解して添加することが好ましい。また、添加された芳香族ジヒドロキシ化合物やジアリールカーボネートは該反応器内でプレポリマーと反応することが必要である。その為、該反応器内でのプレポリマーの滞留時間は１５分以上あることが好ましく、特に３０分以上あることが好ましい。
【００４１】
水酸基比率を制御するために添加されるジアリールカーボネートもしくは芳香族ジヒドロキシ化合物の添加量は、下記式１及び式２とからジアリールカーボネートの量Wbiが、下記式３及び式４から芳香族ジヒドロキシ化合物の量Wciが求められる。その際、求めたWbi及びWciの内、正の値として算出された成分を添加する。但し、連続してジアリールカーボネートもしくは芳香族ジヒドロキシ化合物を添加して水酸基末端比率を制御している場合には、添加成分がジアリールカーボネートから芳香族ジヒドロキシ化合物へ、もしくは芳香族ジヒドロキシ化合物からジアリールカーボネートへ変更になる場合には、それぞれ下記式６、下記式７式から算出される添加量を添加する。
【００４２】
【数１０】

【００４３】
【数１１】

【００４４】
【数１２】

【００４５】
【数１３】

【００４６】
【数１４】

【００４７】
【数１５】

【００４８】
式中、OH₀はプレポリマーの目標水酸基末端比率（モル％）、Mn₀はプレポリマーの目標数平均分子量、OH_aはプレポリマーの実測水酸基末端比率、Mn_aはプレポリマーの実測数平均分子量、Mn_bは添加するジアリールカーボネートの分子量、Mn_cは添加する芳香族ジヒドロキシ化合物の分子量、X_biは時点ｉで添加するジアリールカーボネートのモル分率量、X_ciは時点ｉで添加する芳香族ジヒドロキシ化合物のモル分率、W_aiは時点ｉで反応器を通過するプレポリマーの重量（ｇ/hr）、Wbiは時点ｉで添加するジアリールカーボネートの重量（ｇ/hr）、Wciは時点ｉで添加する芳香族ジヒドロキシ化合物の重量（ｇ/hr）である。aは試行錯誤で求められる装置係数であり、0.5〜1.0の範囲にある。
【００４９】
具体的に例を挙げて上記を説明する。芳香族ジヒドロキシ化合物としてのビスフェノールＡとジアリールカーボネートとしてのジフェニルカーボネートから複数の反応器を用いてポリカーボネートを製造する際に、水酸基末端比率を制御する反応器において、プレポリマーの目標数平均分子量を3800、目標水酸基末端比率を40モル％として連続反応を開始したところ、数平均分子量が3700、水酸基末端比率が44モル％のプレポリマーが製造されていた。反応器内通過するプレポリマー量は10kgであった。目標水酸基末端比率よりも測定値が大きいので、反応器内に添加するジフェニルカーボネートの量（W_bi）を求めた。これら数値及びa=1を上記式１及び２に挿入すると、反応器内に添加するジフェニルカーボネートのモル分率Xbiは0.07、添加重量（W_b）は、43g/hrとなり、同様に、上記式３及び４より芳香族ジヒドロキシ化合物のモル分率Xci及び添加重量Wciを求めると、それぞれ−0.1655、−85ｇ/hrとなる。従って、この時点ではジフェニルカーボネートを43g/hrで添加することで水酸基末端比率を制御できる。
【００５０】
本発明においては、水酸基末端比率を上記のように制御することで、目標水酸基末端比率とポリカーボネートプレポリマーの水酸基末端比率との差を±１０％に制御することが好ましい。更に好ましくは±７％の範囲にあり、特に好ましくは±４％の範囲にある。上記範囲よりも大きい場合には、得られるポリカーボネートの分子量変動幅が増大しやすい傾向にある。プレポリマーの水酸基末端比率測定値と目標水酸基末端比率との差（Ｄ）は、下記式８で定義される。
【００５１】
【数１６】

【００５２】
エステル交換反応によるポリカーボネートの製造は、芳香族モノヒドロキシ化合物を脱離しながら反応が進行するので、分子量が増加するにつれて水酸基末端比率は減少もしくは増加する。その為、ポリカーボネートプレポリマーの目標水酸基末端比率は、目標とするプレポリマーの分子量を基準とする水酸基末端比率となる。プレポリマーの分子量が目標水酸基末端比率の基準分子量から大きくずれていれば、上記式８に代入する水酸基末端比率測定値は分子量補正が必要である。補正は、予め分子量と水酸基末端比率との相関を求めておくことで可能である。また、触媒の種類や量、反応温度、減圧度、装置の大きさや種類等の製造条件によっても、分子量と水酸基末端比率との相関及び目標水酸基末端比率は異なるものとなる。その為、若干の試行錯誤後に経験則によって、相関を求めると共に、目標とする水酸基末端比率は決定される。
【００５３】
上記のようにして水酸基末端比率が制御された工程Ｂのポリカーボネートプレポリマーは、次工程の後重合工程Ｃへ送られる。本発明では、一系列のＡ工程に対して、Ｃ工程は複数ある。一系列のＢ工程に対しては複数のＣ工程があっても良いし、複数のＢ工程に対して同数もしくはそれ以上のＣ工程があってもよい。また、Ｃ工程の内、途中から反応器が並列にある場合、Ｃ工程は複数と見なし、最終反応器の系列の数だけ系列があるものとする。
【００５４】
Ｃ工程で用いられる前述の反応器が全て使用できるが、本工程では溶融粘度が高く、フェノールの脱離が律速となるため、薄膜反応器、遠心式薄膜蒸発反応器、表面更新型二軸混練反応器、二軸横型攪拌反応器、濡れ壁式反応器、自由落下させながら重合する多孔板型反応器、ワイヤーに沿わせて落下させながら重合するワイヤー付き多孔板型反応器等を組み合わせて使用することが好ましい。特に最終反応器には、ワイヤーに沿わせて落下させながら重合するワイヤー付き多孔板型反応器を用いることが好ましい。反応温度は、通常１００〜３５０℃にあり、好ましくは２００〜３２０℃の温度の範囲で選ばれる。より好ましくは２３０〜３００℃、特に好ましくは２５０〜２８０℃の範囲にある。反応圧力は１００ｍｍＨｇ以下、好ましくは５０ｍｍＨｇ以下、より好ましくは２０ｍｍＨｇ以下である。特に最終重合器においては２ｍｍＨｇ以下が好ましい。本発明の方法で得られる芳香族ポリカーボネートの数平均分子量は、通常５０００〜１０００００の範囲であり、好ましくは５０００〜３００００の範囲である。
【００５５】
本発明においては、最終反応器出口以降で該ポリカーボネートの溶融粘度を測定して、最終反応器の減圧度を制御することが好ましい。溶融粘度の測定は、通常使用されているオンライン型の自動粘度計や自動ＭＩ計を用いて連続的に測定することが好ましい。分子量と溶融粘度との間には相関関係があることから、予めこの相関関係を求めておけば、溶融粘度はオンラインで測定できるので分子量を測定するよりも最終反応器の制御には適している。また、溶融粘度の測定誤差や過剰制御の影響を低下させるために、３〜１０測定値の平均値を用いて、最終反応器の減圧度を制御することが好ましい。最終反応器の減圧度制御は、減圧度と溶融粘度との関係を予め数式化しておき、該数式がプログラムされたコンピュータを用いることで、連続的に自動制御することが好ましい。溶融粘度を測定する個所は、最終反応器出口以降であればよいが、反応器出口に近い場所に設置することが測定と制御との間に時間的遅れが少なくなり好ましい。
【００５６】
本発明においては、複数のＣ工程から異なる分子量のポリカーボネートが製造されるが、そのままそれぞれ単独にペレット化しても良いし、混合して分子量や分子量分布の異なるＰＣを製造してもよい。また、安定剤、酸化防止剤、染顔料、紫外線吸収剤、難燃剤等の添加剤や、ガラス繊維、フィラーとの強化剤等を含むポリカーボネート組成物を製造する場合には、各Ｃ工程の最終反応器から溶融状態のままで押出機やスタティックミキサー等にポリカーボネートを供給して、前述の添加剤等を添加・溶融混練してペレット化することが好ましい。また、光学材料用ポリカーボネートを製造する場合には、得られたポリカーボネートの異物を低減させるためにポリマーフィルターを設置することも好ましく行われる。
【００５７】
【発明の実施の形態】
以下、本発明の実施の形態を詳細に説明する。
（１）数平均分子量（以下、Ｍｎと略す）の測定
数平均分子量（以下、Ｍｎと略す）は、ゲルパーミエーションクロマトグラフィー（ＧＰＣ）を用いて行った。テトラヒドロフラン溶媒、ポリスチレンゲルを使用し、標準単分散ポリスチレンの構成曲線から下式による換算分子量較正曲線を用いて求めた。
Ｍ_PC＝０．３５９１Ｍ_PS ^1.0388
（Ｍ_PCはポリカーボネートの分子量、Ｍ_PSはポリスチレンの分子量）
（２）水酸基末端比率の測定
ポリカーボネートプレポリマー0.3gを5mlの重水素置換クロロホルムに溶解し、23℃で1H-NMR（バルカー社製EX-400）を用いて末端基を測定した。水酸基末端比率（モル％）は全末端数に対する水酸基の割合により計算した。
【００５８】
【実施例１】
図１に示すようなプロセスで、分子量と水酸基末端比率の異なる２種類のポリカーボネート（数平均分子量１０６００／水酸基末端比率３０モル％のポリカーボネートと、数平均分子量１２６００／水酸基末端比率１７モル％）を同時に連続製造した。
原料溶解混合槽３ａ、３ｂはバッチ的に交互に運転し、貯槽１０以降は連続的に運転した。原料溶解混合槽３ａ、３ｂ及び撹拌槽型重合器１７、２６ａ、２６ｂは、いずれも攪拌翼を備えており、接液部の材質はＳＵＳ３１６Ｌである。貯槽１０の接液部の材質もＳＵＳ３１６Ｌである。横型二軸攪拌型重合器３５ａ、３５ｂは、Ｌ／Ｄ＝６で回転直径１４０ｍｍの二軸の攪拌羽根を有しており、接液部の材質はＳＵＳ３１６Ｌである。ワイヤ付多孔板型重合器４２ａ、４２ｂは、孔径５ｍｍの孔を５０個有する多孔板４３ａ、４３ｂを備えており、接液部の材質はＳＵＳ３１６Ｌである。孔の中心から鉛直に１ｍｍ径のＳＵＳ３１６Ｌ製ワイヤ４４ａ、４４ｂを重合器下部の液溜まで垂らしてあり、落下する高さは８ｍである。各溶融混合槽及び重合器間をつなぐ移送配管７ａ、７ｂ、１３、１５、２４、３３ａ、３３ｂ、４０ａ、４０ｂ、５０ａ、５０ｂの接液部の材質はいずれもＳＵＳ３１６Ｌである。各混合槽、重合器、貯槽及び全ての移送配管の接液部は重合に先立ちフェノールで１５０℃で洗浄されている。攪拌槽型第１重合器１７までがポリカーボネートプレポリマー製造工程Ａ、攪拌槽型第２重合器２６ａ、２６ｂがポリカーボネートプレポリマーの水酸基末端比率制御工程Ｂ、横型二軸攪拌型重合器３５ａ、３５ｂ以降が後重合工程Ｃに相当する。
【００５９】
原料溶解混合槽３ａ、３ｂは、反応温度１８０℃、反応圧力常圧、シール窒素（酸素濃度０．５ｐｐｍ）ガス流量１リットル／ｈｒの条件である。原料溶解混合槽３ａに、５０ｍｍＨｇで真空窒素置換を５回したビスフェノールＡ粉体とジフェニルカーボネート粉体（対ビスフェノールＡモル比１．１０）を８０Ｋｇと、水酸化ナトリウム７ｍｇを仕込み２．５時間溶融混合し、溶融混合物を全量、移送配管７ａから移送貯槽１０に移送した。溶融混合物はエステル交換反応が進行しており、ビスフェノールＡの反応率は７０％であった。同様に原料溶解混合槽３ｂで同様に原料を溶解混合し、貯槽１０の溶融ポリマーがなくなる前に貯槽１０に移送した。原料溶解混合槽３ａ，３ｂは交互に溶解混合及び移送を繰り返した。貯槽１０は、常圧、１８０℃に保たれている。貯槽１０に移送された溶融ポリマーは、２０ｋｇ／ｈｒで連続に攪拌槽型第１重合器１７に供給した。
【００６０】
攪拌槽型第１重合器１７は、反応温度２３５℃、反応圧力９８ｍｍＨｇの条件であり、溶融ポリマーの液容量が４０リットルに達したら、液容量４０リットルを一定に保つように攪拌槽型第２重合器２６ａ、２６ｂに数平均分子量８５０の溶融ポリマーを等量連続的に供給した。攪拌槽型第２重合器２６ａ、２６ｂは、反応温度２５２℃、反応圧力６ｍｍＨｇの条件であり、溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを一定に保つように横型二軸攪拌型重合器３５ａ、３５ｂに溶融ポリマーを連続に供給した。攪拌槽型第２重合器の出口付近のサンプリング口５２ａ、５２ｂ及び自動粘度計５３ａ、５３ｂでモニタリングしたプレポリマーの水酸基末端比率は４３％、数平均分子量は２５５０であった。攪拌槽型第２重合器２６ａで製造されるポリカーボネートプレポリマーの目標が４１％、２６００であったので、式１及び式２よりＸｂｉ＝０．０３５７、Ｗｂｉ＝０．０１７を求め、０．０１７ｋｇ／ｈｒのジフェニルカーボネートを溶融状態で注入口２８ａから攪拌槽型第２重合器２６ａに供給した。一方、攪拌槽型第２重合器２６ｂで製造されるポリカーボネートプレポリマーの目標は３８％、２６００であったので、式１及び式２より、Ｘｂｉ＝０．０８８８，Ｗｂｉ＝０．０４６を求め、０．０４６ｋｇ／ｈｒのジフェニルカーボネートを溶融状態で注入口２８ｂから攪拌槽型第２重合器２６ｂに供給した。その後、攪拌槽型第２重合器２６ａ、２６ｂ共に、３０分間隔で水酸基末端比率及び分子量を測定し、同様に式１、式２、式３、式４、式６及び式７に基づき算出されたジフェニルカーボネートもしくはビスフェノールＡ量を供給し続けて水酸基末端比率を連続的に制御した。
【００６１】
横型二軸攪拌型重合器３５ａ、３５ｂでは、反応温度２６５℃、反応圧力０．６ｍｍＨｇの条件であり、溶融ポリマーの液容量が１０リットルに達したら、液容量１０リットルを一定に保つようにワイヤ付多孔板型重合器４２ａ、４２ｂにそれぞれ数平均分子量６１００、６０００の溶融ポリマーを供給した。ワイヤ付多孔板型重合器４２ａは、反応温度２６５℃に保持されており、オンライン粘度計５４ａの測定値に基づき、目標数平均分子量１０６００を得るように反応圧力０．４±０．３ｍｍＨｇの範囲で制御した。重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように、移送配管５０ａを経て抜き出し口５１ａよりポリカーボネートを抜き出した。得られたポリカーボネートの数平均分子量は１０６００、水酸基末端比率は３０モル％であった。一方、ワイヤ付多孔板型重合器４２ｂは、反応温度２７２℃に保持されており、オンライン粘度計５４ｂの測定値に基づき、目標数平均分子量１２６００を得るように反応圧力０．２±０．２ｍｍＨｇの範囲で制御した。重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように、移送配管５０ｂを経て抜き出し口５１ｂよりポリカーボネートを抜き出した。得られたポリカーボネートの数平均分子量は１２６００、水酸基末端比率は１７モル％であった。
【００６２】
連続運転は開始から８００時間行った。この間、３時間ごとに測定した２種類のポリカーボネートの分子量と水酸基末端比率の変動は、それぞれ数平均分子量１０６００±３００／水酸基末端比率３０±１モル％、数平均分子量１２６００±４００／水酸基末端比率１７±１モル％の範囲にあり、分子量と水酸基末端比率が異なる２種類のポリカーボネートを切り替えロス及び品質変動もなく安定的に連続製造することができた。得られたポリカーボネートの色調も良好であった。
【００６３】
【比較例１】
図２に示すようなプロセスで、分子量と水酸基末端比率の異なる２種類のポリカーボネートを製造条件を切り替えることで製造した。最初に数平均分子量１０６００／水酸基末端比率３０モル％のポリカーボネートを製造した後、原料の初期仕込みモル比及び最終反応器の温度を変更して数平均分子量１２６００／水酸基末端比率１７モル％のポリカーボネートを製造した。但し、用いた装置は図１と同様であり、２６ａ以降の系列のみを使用した。
【００６４】
原料溶解混合槽３ａ、３ｂは、反応温度１８０℃、反応圧力常圧、シール窒素（酸素濃度０．５ｐｐｍ）ガス流量１リットル／ｈｒの条件である。原料溶解混合槽３ａに、５０ｍｍＨｇで真空窒素置換を５回したビスフェノールＡ粉体とジフェニルカーボネート粉体（対ビスフェノールＡモル比１．１０）を８０Ｋｇと、水酸化ナトリウム７ｍｇを仕込み２．５時間溶融混合し、溶融混合物を全量、移送配管７ａから移送貯槽１０に移送した。溶融混合物はエステル交換反応が進行しており、ビスフェノールＡの反応率は７０％であった。同様に原料溶解混合槽３ｂで同様に原料を溶解混合し、貯槽１０の溶融ポリマーがなくなる前に貯槽１０に移送した。原料溶解混合槽３ａ，３ｂは交互に溶解混合及び移送を繰り返した。貯槽１０は、常圧、１８０℃に保たれている。貯槽１０に移送された溶融ポリマーは、１０ｋｇ／ｈｒで連続に攪拌槽型第１重合器１７に供給した。
【００６５】
攪拌槽型第１重合器１７は、反応温度２３５℃、反応圧力１０５ｍｍＨｇの条件であり、溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを一定に保つように攪拌槽型第２重合器２６ａに数平均分子量８５０の溶融ポリマーを等量連続的に供給した。攪拌槽型第２重合器２６ａは、反応温度２５２℃、反応圧力６ｍｍＨｇの条件であり、溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを一定に保つように横型二軸攪拌型重合器３５ａに溶融ポリマーを連続に供給した。攪拌槽型第２重合器の出口付近のサンプリング口５２ａ及び自動粘度計５３ａでモニタリングしたプレポリマーの水酸基末端比率は４０％、数平均分子量は２６５０であった。
【００６６】
横型二軸攪拌型重合器３５ａでは、反応温度２６５℃、反応圧力０．６ｍｍＨｇの条件であり、溶融ポリマーの液容量が１０リットルに達したら、液容量１０リットルを一定に保つようにワイヤ付多孔板型重合器４２ａに数平均分子量５９００の溶融ポリマーを供給した。ワイヤ付多孔板型重合器４２ａは、反応温度２６５℃に保持されており、オンライン粘度計５４ａの測定値に基づき、目標数平均分子量１０６００を得るように反応圧力０．４±０．３ｍｍＨｇの範囲で制御した。重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように、移送配管５０ａを経て抜き出し口５１ａよりポリカーボネートを抜き出した。得られたポリカーボネートの数平均分子量は９８００、水酸基末端比率は２３モル％であった。目標とするポリカーボネートと比べて水酸基末端比率が低いので、原料仕込みモル比１．１０を１．０８に変更して連続運転を継続した。その後、試行錯誤で２度原料仕込みモル比を変更した後、最終的に原料仕込み１．０９とすることで、数平均分子量１０６００、水酸基末端比率３０％のポリカーボネートを得た。この間、目標と異なる約１００ｋｇのポリカーボネートを製造して損失を被った。また、同条件で連続運転を継続し、３時間ごとに数平均分子量と水酸基末端比率を測定したが、数平均分子量１０６００±３００／水酸基末端比率３０±１％の範囲にあるポリカーボネートの収率は７５％であった。
【００６７】
次に、数平均分子量１２６００／水酸基末端比率１７モル％のポリカーボネートを得るために、仕込みモル比を１．０８、ワイヤ付多孔板型重合器４２ａの温度を２７２℃、圧力を０．２±０．２ｍｍＨｇに変更して製造を続けた。その後仕込みモル比を３回変更して、目標のポリカーボネートを得たが、この間約２５０ｋｇのポリカーボネートの切り替えのためにロスした。その後、８００時間の連続運転を実施し３時間ごとに数平均分子量と水酸基末端比率を測定したが、数平均分子量１２６００±４００／水酸基末端比率１７±１モル％の範囲にあるポリカーボネートは７３％であった。
【００６８】
【実施例２】
図３に示すようなプロセスで、分子量の異なる２種類のポリカーボネート（数平均分子量８５００／水酸基末端比率３２モル％と数平均分子量１３０００／水酸基末端比率２５モル％）を同時に連続製造した。但し、用いた装置３ａ、３ｂ、１０，１７，２６ａ、４２ａ、４２ｂは図１と同様であり、更に４２ａと同型のワイヤ付多孔板型重合器５６、スタティックミキサー６７を使用した。
【００６９】
原料溶解混合槽３ａ、３ｂは、反応温度１８２℃、反応圧力常圧、シール窒素（酸素濃度０．５ｐｐｍ）ガス流量１リットル／ｈｒの条件である。原料溶解混合槽３ａに、５０ｍｍＨｇで真空窒素置換を５回したビスフェノールＡ粉体とジフェニルカーボネート粉体（対ビスフェノールＡモル比１．０８）を８０Ｋｇと、水酸化ナトリウム１２ｍｇを仕込み２．５時間溶融混合し、溶融混合物を全量、移送配管７ａから移送貯槽１０に移送した。溶融混合物はエステル交換反応が進行しており、ビスフェノールＡの反応率は７８％であった。同様に原料溶解混合槽３ｂで同様に原料を溶解混合し、貯槽１０の溶融ポリマーがなくなる前に貯槽１０に移送した。原料溶解混合槽３ａ，３ｂは交互に溶解混合及び移送を繰り返した。貯槽１０は、常圧、１８２℃に保たれている。貯槽１０に移送された溶融ポリマーは、２０ｋｇ／ｈｒで連続に攪拌槽型第１重合器１７に供給した。
【００７０】
攪拌槽型第１重合器１７は、反応温度２３５℃、反応圧力１０５ｍｍＨｇの条件であり、溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを一定に保つように攪拌槽型第２重合器２６ａに数平均分子量８６０の溶融ポリマーを等量連続的に供給した。攪拌槽型第２重合器２６ａは、反応温度２５４℃、反応圧力３ｍｍＨｇの条件であり、溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを一定に保つようにワイヤ付多孔板型重合器５６に数平均分子量が２３３０の溶融ポリマーを連続に供給した。ワイヤ付多孔板型重合器５６では、反応温度２６０℃、反応圧力０．８ｍｍＨｇの条件であり、重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように２６０℃に保たれたスタティックミキサー６７に供給した。サンプリング口６８から分取したポリカーボネートプレポリマー及びインライン粘度計から該溶融ポリマーの数平均分子量は５５００，水酸基末端比率が４１モル％であった。目標の分子量と水酸基末端比率を有するポリカーボネートを得るためにはサンプリング口６８での分析値が数平均分子量５５００，水酸基末端比率が３８モル％であることが必要であるので、式３及び式４を用いて求められたＸｃｉ＝０．０６０５，Ｗｃｉ＝０．０３０から、０．０３０ｋｇ／ｈｒのビスフェノールＡを供給口６６よりスタティックミキサーに供給した。その後、３０分間隔で水酸基末端比率及び分子量を測定し、同様に式１、式２、式３、式４、式６及び式７に基づき算出されたジフェニルカーボネートもしくはビスフェノールＡ量を供給し続けて水酸基末端比率を連続的に制御した。
【００７１】
スタティックミキサーから排出される溶融ポリマーは、ワイヤ付多孔板型重合器４２ａ、４２ｂに等量供給された。ワイヤ付多孔板型重合器４２ａは、反応温度２７６℃に保持されており、オンライン粘度計５４ａの測定値に基づき、数平均分子量１３０００のポリカーボネートを得るように反応圧力を０．３±０．３ｍｍＨｇの範囲で制御した。重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように、移送配管５０ａを経て抜き出し口５１ａよりポリカーボネートを抜き出した。得られたポリカーボネートの数平均分子量は１３０００、水酸基末端比率は２５モル％であった。ワイヤ付多孔板型重合器４２ｂは、反応温度２６０℃に保持されており、オンライン粘度計５４ｂの測定値に基づき、数平均分子量８５００のポリカーボネートを得るように反応圧力を０．６±０．３ｍｍＨｇの範囲で制御した。重合器下部の溶融ポリマーの液容量が２０リットルに達したら、液容量２０リットルを保つように、移送配管５０ｂを経て抜き出し口５１ｂよりポリカーボネートを抜き出した。得られたポリカーボネートの数平均分子量は８５００、水酸基末端比率は３２モル％であった。
【００７２】
連続運転は開始から８００時間行った。この間、３時間ごとに測定した２種類のポリカーボネートの分子量と水酸基末端比率の変動は、それぞれ数平均分子量１３０００±４００／水酸基末端比率２５±１モル％、数平均分子量８５００±３００／水酸基末端比率３２±１モル％の範囲にあり、分子量と水酸基末端比率が異なる２種類のポリカーボネートを切り替えロス及び品質変動もなく安定的に連続製造することができた。得られたポリカーボネートの色調も良好であった。
【００７３】
【発明の効果】
本発明の方法では、実施例で詳細に示したように、得られるポリカーボネートの分子量変動が少なく安定的にポリカーボネートを製造でき、更に分子量や水酸基末端比率の異なる複数のポリカーボネートを製造する際の切り替えロスを減少して且つ一定品質の複数のポリカーボネートを安定的に同時に連続製造できる。
【図面の簡単な説明】
【図１】分子量と水酸基末端比率の異２２種類のポリカーボネートの同時連続製造装置の図である。
【図２】図１と同様で、２６ａ以降の系列のみを使用した装置の図である。
【図３】図１と同様で、更に４２ａと同型の重合器５６とミキサー６７を使用した装置の図である。
【符号の説明】
１ａ，１ｂ 原料供給口
２ａ，２ｂ、９、１９、２８、３６、４６ ベント口
３ａ，３ｂ 原料溶解混合槽
４ａ、４ｂ、１８、２７ａ、２７ｂ、３９ａ、３９ｂ 攪拌軸
５ａ，５ｂ、１１、２０、２９ａ、２９ｂ、４７ａ、４７ｂ、５７ 溶融ポリマー
６ａ．６ｂ、１２、２１、３０ａ、３０ｂ、３７ａ、３７ｂ、４８ａ、４８ｂ、６４ ５２ａ、５２ｂ、６８排出口
７ａ，７ｂ、１３、１５、２４、３３ａ、３３ｂ、４０ａ、４０ｂ、５０ａ、５０ｂ、６５、７０ 移送配管
８、１６、２５ａ、２５ｂ、２８ａ、２８ｂ、３４ａ、３４ｂ、４１ａ、４１ｂ、５５、６６ 供給口
１０ 貯槽
１４、２３、３２ａ、３２ｂ、３８ａ、３８ｂ、４９ａ、４９ｂ、５８ 移送ポンプ
１７ 攪拌槽型第１重合器
２２、３１ａ、３１ｂ、４５ａ、４５ｂ、６３ ガス供給口
２６ａ、２６ｂ 攪拌槽型第２重合器
３５ａ、３５ｂ 横型二軸攪拌型重合器
４２ａ、４２ｂ、５６ ワイヤ付多孔板型重合器
４３ａ、４３ｂ、６１ 多孔板
４４ａ、４４ｂ、６２ ワイヤ
５１ａ、５１ｂ 製品排出口
５３ａ、５３ｂ、５４ａ、５４ｂ インライン粘度計
６７ スタティックミキサー[0001]
BACKGROUND OF THE INVENTION
The present invention relates to an industrial continuous production method of polycarbonate. Specifically, the present invention relates to a continuous production method for producing a plurality of polycarbonates having different molecular weights stably with little variation in molecular weight and efficiently with little loss.
[0002]
[Prior art]
In recent years, aromatic polycarbonates are widely used in many fields as engineering plastics having excellent heat resistance, impact resistance, transparency, and the like, and their importance is increasing.
With respect to the method for producing this aromatic polycarbonate, various polymerization methods have been studied. Among them, an interfacial polycondensation method in which an alkaline aqueous solution of an aromatic dihydroxy compound, for example, 2,2′-bis (4-hydroxyphenyl) propane (hereinafter referred to as bisphenol A) and phosgene is reacted in the presence of an organic solvent is known. It is. The organic solvent used in this method is a halogen-based organic solvent, and for example, methylene chloride, chlorobenzene and the like are used, and in particular, methylene chloride is mainly used. However, in this method, it is difficult to completely remove the organic solvent from the polymer obtained, and mold corrosion or coloring due to the halogen derived from the remaining organic solvent occurs, which adversely affects the final use.
[0003]
Therefore, a method for producing an aromatic polycarbonate from an aromatic dihydroxy compound and diaryl carbonate has been studied. For example, a method of producing polycarbonate by eliminating phenol by transesterification in a molten state of bisphenol A and diphenyl carbonate is known as a so-called transesterification method or melting method. In that case, the melt viscosity varies greatly depending on the molecular weight of the polycarbonate prepolymer in the middle of the reaction, and the higher the molecular weight, the more it is necessary to react at high temperature and high vacuum to promote the elimination of phenol. Is manufactured by connecting a plurality of reactors in series and gradually changing the temperature and the degree of vacuum. For example, in Japanese Patent Laid-Open No. 2-153923, polycarbonate is manufactured by connecting three tank reactors and a total of five reactors of a centrifugal thin film evaporation reactor and a horizontal stirring reactor in series.
[0004]
However, the production of polycarbonate by the transesterification method is difficult to control the reaction compared to the production by the interfacial polycondensation method even if the reaction is controlled in multiple stages as described above. There are disadvantages such as being unable to control the stability stably. For this reason, methods for producing polycarbonates with stable molecular weight control have been studied. For example, Japanese Patent Laid-Open No. 6-65366 proposes a method of controlling a reaction by providing a viscometer at the final reactor outlet and feeding back the viscosity measurement result to the reaction conditions of the final reactor. However, in this method, it is difficult to completely control the reaction system in the polycarbonate prepolymer stage having large fluctuations, and the molecular weight of the obtained polycarbonate varies greatly due to the quality fluctuation of the prepolymer, and can be stably controlled. There wasn't. Japanese Patent Application Laid-Open No. 8-113737 proposes a method for stably producing by continuously measuring the amount of phenols produced during the production of the polycarbonate prepolymer. However, the molecular weight of the polycarbonate obtained by this method could not be controlled stably. Therefore, the advent of a method for stably controlling the molecular weight of the obtained polycarbonate has been desired.
[0005]
Moreover, when manufacturing a polycarbonate industrially, since manufacturing many brands is requested | required, it is necessary to manufacture several polycarbonate from which molecular weight and a hydroxyl-terminal ratio differ. Therefore, it is a common practice to change the reactor temperature and the degree of pressure reduction by adjusting the charged molar ratio of the starting aromatic dihydroxy compound and diaryl carbonate. However, as described above, the industrial production method of polycarbonate by the transesterification method is performed in a production process in which a plurality of reactors are connected in series, so that the charged molar ratio of starting materials and reaction conditions are changed. However, since it takes a lot of time to reach the final reactor, there is a problem that a lot of switching loss occurs until the target polycarbonate is obtained.
[0006]
[Problems to be solved by the invention]
The present invention provides a continuous production method in which the molecular weight fluctuation of the obtained polycarbonate is small, and a continuous production of a polycarbonate of a constant quality stably at the same time with reduced switching loss when producing a plurality of polycarbonates having different molecular weights and hydroxyl end ratios. It aims to provide a way to do that.
[0007]
[Means for Solving the Problems]
In order to solve the above problems, the present inventors have conducted intensive research on the relationship between the prepolymer and the finally obtained polycarbonate by carrying out continuous long-term production of polycarbonate. As a result, the inventors have found the surprising fact that the cause of the molecular weight variation of the obtained polycarbonate is the variation in the terminal ratio of the polycarbonate prepolymer, which has not been noticed at all. Based on such knowledge, the present inventors have further researched and can solve the above problems by controlling the hydroxyl group terminal ratio of the prepolymer, and further by combining with a plurality of post-polymerization steps installed in parallel. The present inventors have found that a plurality of polycarbonates having different molecular weights and hydroxyl terminal ends can be simultaneously stably produced, thereby reaching the present invention.
[0008]
  That is, the present invention
(1) In a method of continuously producing a polycarbonate having a number average molecular weight of 5000 to 30000 by polymerizing in a molten state from an aromatic dihydroxy compound and diaryl carbonate using at least two reactors, a reaction temperature of 150 to 290 Polycarbonate prepolymer production step A having a reaction pressure of 10 mmHg to normal pressure, reaction temperature of 180 to 270 ° C, reaction pressure of 10 mmHg to hydroxyl group terminal ratio control step B of the polycarbonate prepolymer having a reaction pressure of 230 to 300 ° C, It consists of a post-polymerization step C having a reaction pressure of 100 mmHg or less.B processOr multipleProvided in parallelProcess BAre connected and a single B process is provided, a plurality of C processes provided in parallel are connected after the B process, and a plurality of B processes provided in parallel are provided. Later it was provided in a single C process or multiple parallelStep C is connected, and hydroxyl group terminal ratio control step B measures the molecular weight and hydroxyl group terminal ratio of the polycarbonate prepolymer. Based on the results, aromatic dihydroxy compound or diaryl carbonate is added to the polycarbonate prepolymer liquid. The difference between the target hydroxyl group terminal ratio and the hydroxyl group terminal ratio of the polycarbonate prepolymer is controlled within a range of ± 10% by adding to and reacting with the initial raw material by controlling the degree of vacuum in the C process. A method for continuously producing a polycarbonate, comprising simultaneously producing a plurality of polycarbonates having different molecular weights without changing the charged molar ratio of the aromatic dihydroxy compound and diaryl carbonate used as the above.
[0009]
(2) The aromatic dihydroxy compound or diaryl carbonate to be added is added in the amount Wci of the aromatic dihydroxy compound or the amount Wbi of the diaryl carbonate obtained from the following formula 1, formula 2, formula 3, and formula 4. A method for continuously producing a plurality of polycarbonates having different molecular weights as described in (1) above. (However, among the obtained Wbi and Wci, a component calculated as a positive value is added).
[0010]
[Equation 5]

[0011]
[Formula 6]

[0012]
[Expression 7]

[0013]
[Equation 8]

[0014]
(However, in the formula, OH₀Is the target hydroxyl group end ratio (mol%) of the prepolymer, Mn₀Is the target number average molecular weight of the prepolymer, OH_aIs the measured hydroxyl terminal ratio of the prepolymer, Mn_aIs the measured number average molecular weight of the prepolymer, Mn_bIs the molecular weight of the added diaryl carbonate, Mn_cIs the molecular weight of the aromatic dihydroxy compound to be added, X_biIs the molar fraction of diaryl carbonate added at time i, X_ciIs the mole fraction of aromatic dihydroxy compound added at time i, W_aiIs the weight of prepolymer passing through the reactor at time i (g / hr), Wbi is the weight of diaryl carbonate added at time i (g / hr), Wci is the weight of aromatic dihydroxy compound added at time i ( g / hr). a is a device coefficient obtained by trial and error, and is in the range of 0.5 to 1.0. )
Is to provide.
[0015]
The present invention is described in detail below.
In the present invention, the aromatic dihydroxy compound is a compound represented by HO—Ar—OH (wherein Ar represents a divalent aromatic group).
The aromatic group Ar is preferably -Ar, for example.¹-Y-Ar²-A divalent aromatic group represented by the formula (wherein Ar¹And Ar²Each independently represents a divalent carbocyclic or heterocyclic aromatic group having 5 to 70 carbon atoms, and Y represents a divalent alkane group having 1 to 30 carbon atoms. ).
[0016]
Divalent aromatic group Ar¹, Ar²In which one or more hydrogen atoms do not adversely influence the reaction, for example, a halogen atom, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a phenyl group, a phenoxy group, It may be substituted with a vinyl group, a cyano group, an ester group, an amide group, a nitro group, or the like.
Preferable specific examples of the heterocyclic aromatic group include aromatic groups having one or more ring-forming nitrogen atoms, oxygen atoms or sulfur atoms.
Divalent aromatic group Ar¹, Ar²Represents, for example, a group such as substituted or unsubstituted phenylene, substituted or unsubstituted biphenylene, substituted or unsubstituted pyridylene. The substituents here are as described above.
The divalent alkane group Y is, for example, an organic group represented by the following chemical formula 1.
[0017]
[Chemical 1]

[0018]
(Wherein R¹, R², R^Three, R^FourAre each independently hydrogen, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a cycloalkyl group having 5 to 10 ring carbon atoms, and a carbocyclic aromatic having 5 to 10 ring carbon atoms. Group, a C6-C10 carbocyclic aralkyl group. k represents an integer of 3 to 11, and R^FiveAnd R⁶Are individually selected for each X and, independently of one another, represent hydrogen or an alkyl group having 1 to 6 carbon atoms, where X represents carbon. R¹, R², R^Three, R^Four, R^Five, R⁶In the above, in the range where one or more hydrogen atoms do not adversely affect the reaction, for example, a halogen atom, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a phenyl group, a phenoxy group, It may be substituted with a vinyl group, a cyano group, an ester group, an amide group, a nitro group, or the like. )
Examples of such a divalent aromatic group Ar include those represented by the following chemical formula 2.
[0019]
[Chemical formula 2]

[0020]
(Wherein R⁷, R⁸Are each independently a hydrogen atom, a halogen atom, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a cycloalkyl group having 5 to 10 ring carbon atoms, or a phenyl group, and m and n Is an integer from 1 to 4, and when m is 2 to 4, each R⁷May be the same or different, and when n is 2 to 4, each R⁸May be the same or different. )
Further, the divalent aromatic group Ar is -Ar¹-Z-Ar²It may be indicated by −.
(Wherein Ar¹, Ar²Is as described above, Z is a single bond or -O-, -CO-, -S-, -SO.₂-, -SO-, -COO-, -CON (R¹Represents a divalent group such as-. However, R¹Is as described above. )
Examples of such divalent aromatic group Ar include those represented by the following chemical formula 3.
[0021]
[Chemical 3]

[0022]
(Wherein R⁷, R⁸, M and n are as described above. )
Furthermore, specific examples of the divalent aromatic group Ar include substituted or unsubstituted phenylene, substituted or unsubstituted naphthylene, substituted or unsubstituted pyridylene, and the like. The substituents here are as described above.
The aromatic dihydroxy compound used in the present invention may be a single type or two or more types. A typical example of the aromatic dihydroxy compound is bisphenol A.
The diaryl carbonate used in the present invention is represented by the following chemical formula 4.
[0023]
[Formula 4]

[0024]
(Wherein Ar^Three, Ar^FourEach represents a monovalent aromatic group. )
Ar^ThreeAnd Ar^FourRepresents a monovalent carbocyclic or heterocyclic aromatic group, this Ar^Three, Ar^FourIn which one or more hydrogen atoms do not adversely influence the reaction, for example, a halogen atom, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a phenyl group, a phenoxy group, It may be substituted with a vinyl group, a cyano group, an ester group, an amide group, a nitro group, or the like. Ar^ThreeAnd Ar^FourMay be the same or different.
Monovalent aromatic group Ar^ThreeAnd Ar^FourAs typical examples of these, there can be mentioned a phenyl group, a naphthyl group, a biphenyl group, and a pyridyl group. These may be substituted with one or more substituents as described above.
Preferred Ar^ThreeAnd Ar^FourAs, for example, the following chemical formula 5 and the like can be mentioned.
[0025]
[Chemical formula 5]

[0026]
Representative examples of diaryl carbonates include substituted or unsubstituted diphenyl carbonates represented by the following chemical formula (6).
[0027]
[Chemical 6]

[0028]
(Wherein R⁹And R^TenEach independently represents a hydrogen atom, an alkyl group having 1 to 10 carbon atoms, an alkoxy group having 1 to 10 carbon atoms, a cycloalkyl group having 5 to 10 ring carbon atoms, or a phenyl group, and p and q are 1 When R is an integer of ˜5 and p is 2 or more, each R⁹May be different from each other, and when q is 2 or more, each R^TenMay be different from each other. )
Among these diphenyl carbonates, preferred are symmetric diaryl carbonates such as unsubstituted diphenyl carbonate, lower alkyl-substituted diphenyl carbonates such as ditolyl carbonate and di-t-butylphenyl carbonate, and particularly diaryls having the simplest structure. Unsubstituted diphenyl carbonate, which is carbonate, is preferred.
[0029]
These diaryl carbonates may be used alone or in combination of two or more.
The use ratio (charge ratio) of the aromatic dihydroxy compound and diaryl carbonate varies depending on the molecular weight and terminal group ratio of the target polycarbonate, the type of aromatic dihydroxy compound and diaryl carbonate used, the polymerization temperature, and other polymerization conditions. The diaryl carbonate is usually used in a proportion of 0.9 to 2.5 mol, preferably 0.95 to 2.0 mol, more preferably 0.98 to 1.5 mol, relative to 1 mol of the aromatic dihydroxy compound. . A method is generally used in which the charging ratio of the aromatic dihydroxy compound and the diaryl carbonate is changed depending on the target molecular weight of the polycarbonate to be produced and the hydroxyl group terminal ratio. However, when the charging ratio is changed, it is not preferable because a lot of switching loss is involved until a polycarbonate having a target molecular weight or hydroxyl group terminal ratio is finally obtained. In the present invention, even if the target molecular weight and the hydroxyl group terminal ratio of the polycarbonate to be produced are changed, it is preferable to produce the polycarbonate without changing the charging ratio of the both.
[0030]
The melt polycondensation reaction can be carried out without adding a catalyst, but is carried out in the presence of a catalyst as necessary in order to increase the polymerization rate. The polymerization catalyst is not particularly limited as long as it is used in this field, but alkali metal or alkaline earth metal hydroxides such as lithium hydroxide, sodium hydroxide, potassium hydroxide, and calcium hydroxide. Alkali metal salts, alkaline earth metal salts, quaternary ammonium salts of hydrides of boron and aluminum such as lithium aluminum hydride, sodium borohydride, tetramethylammonium borohydride; lithium hydride, sodium hydride Alkali metal or alkaline earth metal hydrides such as calcium hydride; Alkali metal or alkaline earth metal alkoxides such as lithium methoxide, sodium ethoxide, calcium methoxide; lithium phenoxide, sodium phenoxide, magnesium phenoxide Alkali metal or alkaline earth metal aryloxides such as xoxide, LiO—Ar—OLi, NaO—Ar—ONa (Ar is an aryl group); Alkali metal or alkaline earth metal such as lithium acetate, calcium acetate, sodium benzoate Organic acid salts of: zinc compounds such as zinc oxide, zinc acetate, zinc phenoxide; boron oxide, boric acid, sodium borate, trimethyl borate, tributyl borate, triphenyl borate, (R¹R²R^ThreeR^Four) NB (R¹R²R^ThreeR^Four) Ammonium borates represented by¹R²R^ThreeR^Four) PB (R¹R²R^ThreeR^FourPhosphonium borates (R)¹, R², R^Three, R^FourIs as described in Chemical Formula 1 above. Boron compounds such as silicon oxide, sodium silicate, tetraalkyl silicon, tetraaryl silicon, diphenyl-ethyl-ethoxy silicon, etc .; germanium oxide, germanium tetrachloride, germanium ethoxide, germanium phenoxide, etc. Germanium compounds; tin compounds such as tin oxides, dialkyltin oxides, dialkyltin carboxylates, tin acetates, ethyltin tributoxide and other alkoxy groups or aryloxy groups; organotin compounds; lead oxides Lead compounds such as lead acetate, lead carbonate, basic carbonate, lead and organic lead alkoxides or aryloxides; Onium compounds such as quaternary ammonium salts, quaternary phosphonium salts, quaternary arsonium salts; oxidation A Antimony compounds such as Timon and antimony acetate; Manganese compounds such as manganese acetate, manganese carbonate, and manganese borate; Titanium compounds such as titanium oxide, titanium alkoxide or aryloxide; Zirconium acetate, Zirconium oxide, Zirconium Examples include catalysts such as alkoxides, aryloxides, zirconium compounds such as zirconium acetylacetone.
[0031]
When using a catalyst, these catalysts may be used alone or in combination of two or more. The amount of these catalysts used is usually 10 with respect to 100 parts by weight of the starting aromatic dihydroxy compound.^-8~ 1 part by weight, preferably 10^-7-10^-1It is selected in the range of parts by weight.
The method of the present invention is a production method comprising a polycarbonate prepolymer production step A, a polycarbonate prepolymer hydroxyl group terminal ratio control step B, and a post-polymerization step C, and a plurality of B steps and / or a series of A steps. It is characterized by comprising a plurality of C steps. As the reactor in each step, those widely used as a polycarbonate reactor can be used. For example, a stirred tank reactor, a thin film reactor, a centrifugal thin film evaporation reactor, a surface renewal type twin-screw kneading reactor A biaxial horizontal stirring reactor, a wet wall reactor, a perforated plate reactor that polymerizes while being dropped freely, a perforated plate reactor with a wire that polymerizes while dropping along a wire, and the like are used. The material of these reactors is not particularly limited, but a material containing 20% or more of iron is preferable, and SUS304, SUS316, and SUS316L are particularly preferably used.
[0032]
Polycarbonate prepolymer production process A consists of a raw material metering and supply process such as aromatic dihydroxy compounds and diaryl carbonate, a raw material dissolving tank and a reactor for dissolving them, and a filter is installed after the raw material dissolving tank and / or the reactor. You may do it. The raw material dissolution tank may be performed in a batch manner or continuously, but the former is preferable. At that time, it is preferable to use a plurality of raw material dissolution tanks while switching them in parallel. The reactor may be a single reactor, or a plurality of reactors may be connected in series. Of these former dissolution tanks and reactors, the former, in particular, a stirred tank reactor is preferably used.
[0033]
The raw material aromatic dihydroxy compound and diaryl carbonate are transferred to the raw material dissolution and mixing tank after measurement. The aromatic dihydroxy compound and diaryl carbonate may be previously vacuum-substituted using an inert gas such as nitrogen. This is true. Also preferred is a method in which either one or both of the aromatic dihydroxy compound and diaryl carbonate are transferred to the raw material dissolution and mixing tank in a molten state from which oxygen is removed. In particular, when manufacturing plants for aromatic dihydroxy compounds or diaryl carbonates are adjacent to each other, it is preferable to connect them with piping or the like and directly transfer them to the raw material dissolution and mixing tank. Although the temperature of a raw material dissolution tank is 115-220 degreeC, Preferably it is 120-200 degreeC, More preferably, it exists in the range of 140-190 degreeC. In the raw material dissolution and mixing tank, the aromatic dihydroxy compound and the diaryl carbonate may not only be dissolved and mixed, but also a transesterification reaction may be allowed to proceed. The reaction rate is adjusted in a range of 5 to 95%, preferably adjusted in a range of 30 to 92%, particularly preferably in a range of 50 to 90%, and then transferred to a subsequent step, for example, a storage tank or a reactor.
[0034]
The reaction temperature in the polycarbonate prepolymer production step A is usually selected in the range of 100 to 350 ° C, preferably 150 to 290 ° C. Preferably it exists in the range of 180-270 degreeC, Most preferably, it is the range of 200-250 degreeC. The number average molecular weight of the obtained polycarbonate prepolymer is in the range of 1500 to 7500. A preferred number average molecular weight range is from 1800 to 6000, more preferably from 2000 to 5000. As the reaction proceeds, an aromatic monohydroxy compound is produced, and the reaction rate can be increased by removing this from the reaction system. Therefore, an inert gas that does not adversely influence the reaction, such as nitrogen, argon, helium, carbon dioxide, and lower hydrocarbon gas, is introduced, and the generated aromatic monohydroxy compound is accompanied by these gases and removed. And a method of reacting under reduced pressure are preferably used. A preferable reaction pressure is in the range of 10 mmHg to normal pressure in the initial stage of polymerization.
[0035]
The prepolymer produced in the polycarbonate prepolymer production step A is sent to the hydroxyl group terminal ratio control step B of the polycarbonate prepolymer. Step B is a step of controlling the hydroxyl group terminal ratio by measuring the molecular weight and the hydroxyl group terminal ratio of the polycarbonate prepolymer, and adding and reacting an aromatic dihydroxy compound or diaryl carbonate. The process B may be one series or a plurality of series. When a polycarbonate having a plurality of hydroxyl group terminal ratios having completely different relationships between the molecular weight and the hydroxyl group terminal ratio is continuously produced simultaneously, a plurality of corresponding series are required. The reactor in Step B is 1 to 2 reactors per series, and a stirred tank reactor is preferably used in order to sufficiently react with the aromatic dihydroxy compound or diaryl carbonate to be added. The reaction temperature is usually selected in the range of 100 to 350 ° C, preferably 150 to 290 ° C. Preferably it exists in the range of 180-270 degreeC, Most preferably, it is the range of 200-250 degreeC. A preferred reaction pressure is in the range of 10 mmHg to normal pressure.
[0036]
In the present invention, the hydroxyl group terminal ratio of the polycarbonate prepolymer is generally in the range of 5 to 95 mol%, and the target hydroxyl group terminal ratio in this range is controlled. When producing a polycarbonate having a hydroxyl group terminal ratio of 50% or more for the purpose of polymer alloy or the like, the target hydroxyl terminal ratio in the prepolymer stage is 50 to 95 mol%. When producing a polycarbonate having a hydroxyl group terminal ratio of 50% or less suitable for optical use, general injection molding, and extrusion molding, the target hydroxyl terminal ratio in the prepolymer stage is 5 to 50 mol%. In the latter case, the hydroxyl group terminal ratio of the polycarbonate prepolymer is preferably set in the range of 5 to 50 mol%. 20-48 mol% is especially preferable.
[0037]
The molecular weight of the polycarbonate prepolymer extracted from the reactor can be measured by, for example, solution viscosity, GPC, melt viscosity and the like. In general, the solution viscosity method can measure the molecular weight simply and quickly, but since the molecular weight obtained is the viscosity average molecular weight, it is necessary to obtain the correlation between the number average molecular weight and the viscosity average molecular weight in advance. Further, in the melt viscosity method, for example, it can be continuously and automatically measured by installing a viscometer in a pipe or sampling port. At that time, it is necessary to obtain a correlation between the melt viscosity and the number average molecular weight. The higher the measurement frequency, the smaller the molecular weight fluctuation of the polycarbonate obtained, which is preferable. However, the measurement frequency is also limited by the molecular weight measurement method. In general, the interval is several seconds to 6 hours.
[0038]
The hydroxyl group terminal ratio of the polycarbonate prepolymer extracted from the reactor can be measured by, for example, NMR, IR, colorimetric (titanium method), liquid chromatography, or the like. Among these, NMR method and liquid chromatography are preferred because of their rapidness and high accuracy. The higher the measurement frequency, the smaller the fluctuation of the hydroxyl group terminal ratio of the resulting polycarbonate, which is preferable. However, the measurement frequency is also limited by the method for measuring the hydroxyl group terminal ratio. In general, the interval is several seconds to 6 hours. In the production method of the present invention, the terminal of the polycarbonate has a hydroxyl group terminal (A) derived from the starting aromatic dihydroxy compound and an aryl group terminal (B) derived from the starting diaryl carbonate. The hydroxyl group terminal ratio (mol%) is obtained by the following formula 5.
[0039]
[Equation 9]

[0040]
For example, when bisphenol-A and diphenyl carbonate are used as starting materials, the hydroxyl terminal ratio is determined from the number of moles of the hydroxyl terminal (A) and phenyl terminal (B) of bisphenol-A.
In the present invention, in order to control the target hydroxyl group terminal ratio of the polycarbonate prepolymer, a diaryl carbonate or an aromatic dihydroxy compound is added to the prepolymer and reacted. As the aromatic dihydroxy compound or diaryl carbonate added to control the hydroxyl group terminal ratio of the prepolymer, it is preferable to use the same aromatic dihydroxy compound or diaryl carbonate as the starting material. However, other aromatic dihydroxy compounds and diaryl carbonates, polycarbonate polymers and oligomers having aryl carbonate ends at both ends, and polycarbonate polymers and oligomers having hydroxyl ends at both ends can also be used. Moreover, since these additions are performed continuously, it is preferable to use a pump as the addition method, and it is particularly preferable to supply the prepolymer liquid. These compounds can be added even in a solid state, but it is preferable to add them by dissolving them in a solvent which does not affect the molten state or reaction using a pump. Further, it is necessary that the added aromatic dihydroxy compound or diaryl carbonate reacts with the prepolymer in the reactor. Therefore, the residence time of the prepolymer in the reactor is preferably 15 minutes or more, particularly preferably 30 minutes or more.
[0041]
The amount of diaryl carbonate or aromatic dihydroxy compound added to control the hydroxyl ratio is the amount of diaryl carbonate Wbi from formula 1 and formula 2 below, and the amount of aromatic dihydroxy compound from formula 3 and formula 4 below. Wci is required. At that time, a component calculated as a positive value among the obtained Wbi and Wci is added. However, when dihydroxy carbonate or aromatic dihydroxy compound is continuously added to control the hydroxyl group terminal ratio, the added component is changed from diaryl carbonate to aromatic dihydroxy compound, or from aromatic dihydroxy compound to diaryl carbonate. When it becomes, the addition amount calculated from the following formula 6 and the following formula 7 is added.
[0042]
[Expression 10]

[0043]
## EQU11 ##

[0044]
[Expression 12]

[0045]
[Formula 13]

[0046]
[Expression 14]

[0047]
[Expression 15]

[0048]
Where OH₀Is the target hydroxyl group end ratio (mol%) of the prepolymer, Mn₀Is the target number average molecular weight of the prepolymer, OH_aIs the measured hydroxyl terminal ratio of the prepolymer, Mn_aIs the measured number average molecular weight of the prepolymer, Mn_bIs the molecular weight of the added diaryl carbonate, Mn_cIs the molecular weight of the aromatic dihydroxy compound to be added, X_biIs the molar fraction of diaryl carbonate added at time i, X_ciIs the mole fraction of aromatic dihydroxy compound added at time i, W_aiIs the weight of prepolymer passing through the reactor at time i (g / hr), Wbi is the weight of diaryl carbonate added at time i (g / hr), Wci is the weight of aromatic dihydroxy compound added at time i ( g / hr). a is a device coefficient obtained by trial and error, and is in the range of 0.5 to 1.0.
[0049]
The above will be described with specific examples. When a polycarbonate is produced from bisphenol A as an aromatic dihydroxy compound and diphenyl carbonate as a diaryl carbonate using a plurality of reactors, the target number average molecular weight of the prepolymer is 3800 in a reactor for controlling the hydroxyl group terminal ratio. When a continuous reaction was started at a target hydroxyl group terminal ratio of 40 mol%, a prepolymer having a number average molecular weight of 3700 and a hydroxyl group terminal ratio of 44 mol% was produced. The amount of prepolymer passing through the reactor was 10 kg. Since the measured value is larger than the target hydroxyl terminal ratio, the amount of diphenyl carbonate added to the reactor (W_bi) When these numerical values and a = 1 are inserted into the above formulas 1 and 2, the molar fraction Xbi of diphenyl carbonate added to the reactor is 0.07, and the added weight (W_b) Is 43 g / hr. Similarly, when the molar fraction Xci and the added weight Wci of the aromatic dihydroxy compound are obtained from the above formulas 3 and 4, they are −0.1655 and −85 g / hr, respectively. Therefore, at this time, the hydroxyl terminal ratio can be controlled by adding diphenyl carbonate at 43 g / hr.
[0050]
In the present invention, it is preferable to control the difference between the target hydroxyl group terminal ratio and the hydroxyl group terminal ratio of the polycarbonate prepolymer to ± 10% by controlling the hydroxyl group terminal ratio as described above. More preferably, it is in the range of ± 7%, and particularly preferably in the range of ± 4%. When it is larger than the above range, the molecular weight fluctuation range of the obtained polycarbonate tends to increase. The difference (D) between the measured value of the hydroxyl group terminal ratio of the prepolymer and the target hydroxyl group terminal ratio is defined by the following formula 8.
[0051]
[Expression 16]

[0052]
In the production of polycarbonate by transesterification, the reaction proceeds while eliminating the aromatic monohydroxy compound, so that the hydroxyl terminal ratio decreases or increases as the molecular weight increases. Therefore, the target hydroxyl group terminal ratio of the polycarbonate prepolymer is a hydroxyl group terminal ratio based on the molecular weight of the target prepolymer. If the molecular weight of the prepolymer is greatly deviated from the reference molecular weight of the target hydroxyl group terminal ratio, the hydroxyl group terminal ratio measurement value substituted into the above formula 8 needs to be corrected for molecular weight. Correction can be made by obtaining a correlation between the molecular weight and the hydroxyl group terminal ratio in advance. Further, the correlation between the molecular weight and the hydroxyl group terminal ratio and the target hydroxyl group terminal ratio vary depending on the production conditions such as the type and amount of the catalyst, the reaction temperature, the degree of reduced pressure, the size and type of the apparatus, and the like. Therefore, a correlation is obtained by a rule of thumb after some trial and error, and the target hydroxyl terminal ratio is determined.
[0053]
The polycarbonate prepolymer of Step B in which the hydroxyl terminal ratio is controlled as described above is sent to the post-polymerization step C of the next step. In the present invention, there are a plurality of C processes for one series of A processes. There may be a plurality of C steps for a series of B steps, and there may be the same or more C steps for a plurality of B steps. In addition, when there are reactors in parallel from the middle of the C process, it is assumed that there are a plurality of C processes, and there are as many series as there are series of final reactors.
[0054]
All the above-mentioned reactors used in Step C can be used, but in this step, the melt viscosity is high and the elimination of phenol is rate-limiting, so a thin film reactor, a centrifugal thin film evaporation reactor, a surface renewal type biaxial kneading Used in combination with reactor, biaxial horizontal stirring reactor, wet wall reactor, perforated plate reactor that polymerizes while dropping freely, perforated plate reactor with wire that polymerizes while dropping along wire, etc. It is preferable to do. In particular, it is preferable to use a perforated plate reactor with a wire that polymerizes while being dropped along the wire for the final reactor. The reaction temperature is usually 100 to 350 ° C., preferably 200 to 320 ° C. More preferably, it is 230-300 degreeC, Most preferably, it exists in the range of 250-280 degreeC. The reaction pressure is 100 mmHg or less, preferably 50 mmHg or less, more preferably 20 mmHg or less. Particularly in the final polymerization vessel, 2 mmHg or less is preferable. The number average molecular weight of the aromatic polycarbonate obtained by the method of the present invention is usually in the range of 5,000 to 100,000, preferably 5,000 to 30,000.
[0055]
In the present invention, it is preferable to control the degree of vacuum of the final reactor by measuring the melt viscosity of the polycarbonate after the outlet of the final reactor. The melt viscosity is preferably measured continuously using a commonly used online type automatic viscometer or automatic MI meter. Since there is a correlation between molecular weight and melt viscosity, if this correlation is obtained in advance, melt viscosity can be measured online, so it is more suitable for controlling the final reactor than measuring molecular weight. . In order to reduce the measurement error of melt viscosity and the influence of excessive control, it is preferable to control the degree of vacuum in the final reactor using an average value of 3 to 10 measurement values. The pressure reduction control of the final reactor is preferably continuously and automatically controlled by formulating the relationship between the pressure reduction degree and the melt viscosity in advance and using a computer programmed with the mathematical expression. The location for measuring the melt viscosity may be after the final reactor outlet, but it is preferable to install the melt viscosity at a location near the reactor outlet because the time delay between measurement and control is reduced.
[0056]
In the present invention, polycarbonates having different molecular weights are produced from a plurality of C steps, but they may be individually pelletized as they are, or may be mixed to produce PCs having different molecular weights and molecular weight distributions. In the case of producing a polycarbonate composition containing additives such as stabilizers, antioxidants, dyes and pigments, ultraviolet absorbers, flame retardants, reinforcing agents with glass fibers and fillers, etc. It is preferable to supply polycarbonate to an extruder, a static mixer or the like in the molten state from the reactor, add the above-mentioned additives, etc., melt knead and pelletize. Moreover, when manufacturing the polycarbonate for optical materials, in order to reduce the foreign material of the obtained polycarbonate, installing a polymer filter is also performed preferably.
[0057]
DETAILED DESCRIPTION OF THE INVENTION
Hereinafter, embodiments of the present invention will be described in detail.
(1) Measurement of number average molecular weight (hereinafter abbreviated as Mn)
The number average molecular weight (hereinafter abbreviated as Mn) was performed using gel permeation chromatography (GPC). A tetrahydrofuran solvent and polystyrene gel were used, and the molecular weight was determined from the standard monodisperse polystyrene composition curve using a reduced molecular weight calibration curve according to the following formula.
M_PC= 0.3591M_PS ^1.0388
(M_PCIs the molecular weight of polycarbonate, M_PSIs the molecular weight of polystyrene)
(2) Measurement of hydroxyl terminal ratio
Polycarbonate prepolymer (0.3 g) was dissolved in 5 ml of deuterium-substituted chloroform, and the end group was measured at 23 ° C. using 1H-NMR (EX-400 manufactured by VALQUA). The hydroxyl terminal ratio (mol%) was calculated by the ratio of hydroxyl groups to the total number of terminals.
[0058]
[Example 1]
In the process shown in FIG. 1, two types of polycarbonates having different molecular weights and hydroxyl terminal ratios (number average molecular weight 10600 / hydroxyl terminal ratio 30 mol% polycarbonate and number average molecular weight 12600 / hydroxyl terminal ratio 17 mol%) are simultaneously applied. Continuous production.
The raw material dissolution and mixing tanks 3a and 3b were alternately operated in batches, and the storage tank 10 and subsequent tanks were continuously operated. The raw material dissolution and mixing tanks 3a and 3b and the stirring tank type polymerizers 17, 26a and 26b are all provided with stirring blades, and the material of the liquid contact part is SUS316L. The material of the liquid contact part of the storage tank 10 is also SUS316L. The horizontal biaxial agitation polymerizers 35a and 35b have biaxial agitation blades having a rotation diameter of 140 mm and L / D = 6, and the material of the liquid contact part is SUS316L. The perforated plate polymerizers 42a and 42b with wires include perforated plates 43a and 43b having 50 holes with a diameter of 5 mm, and the material of the liquid contact part is SUS316L. The SUS316L wires 44a and 44b having a diameter of 1 mm are suspended vertically from the center of the hole to the liquid reservoir at the bottom of the polymerization vessel, and the height of dropping is 8 m. The material of the wetted parts of the transfer pipes 7a, 7b, 13, 15, 24, 33a, 33b, 40a, 40b, 50a, 50b connecting the respective melt mixing tanks and the polymerization vessel is SUS316L. The liquid contact parts of each mixing tank, polymerization vessel, storage tank and all transfer pipes are washed with phenol at 150 ° C. prior to polymerization. Up to the stirring tank type first polymerizer 17 is the polycarbonate prepolymer production step A, the stirring tank type second polymerizers 26a and 26b are the hydroxyl group terminal ratio control step B of the polycarbonate prepolymer, the horizontal biaxial stirring type polymerizers 35a and 35b and thereafter. Corresponds to post-polymerization step C.
[0059]
The raw material dissolution and mixing tanks 3a and 3b are under the conditions of a reaction temperature of 180 ° C., a reaction pressure and a normal pressure, and a seal nitrogen (oxygen concentration of 0.5 ppm) gas flow rate of 1 liter / hr. 80 kg of bisphenol A powder and diphenyl carbonate powder (with a molar ratio of bisphenol A of 1.10) of 50 mmHg and vacuum nitrogen substitution 5 times were charged in the raw material dissolution mixing tank 3a and 7 mg of sodium hydroxide were melted for 2.5 hours. After mixing, the entire amount of the molten mixture was transferred from the transfer pipe 7a to the transfer storage tank 10. The melt mixture had undergone a transesterification reaction, and the reaction rate of bisphenol A was 70%. Similarly, the raw materials were similarly dissolved and mixed in the raw material dissolution and mixing tank 3b, and transferred to the storage tank 10 before the molten polymer in the storage tank 10 disappeared. The raw material dissolution and mixing tanks 3a and 3b repeated dissolution and mixing and transfer alternately. The storage tank 10 is maintained at normal pressure and 180 ° C. The molten polymer transferred to the storage tank 10 was continuously supplied to the stirring tank type first polymerization device 17 at 20 kg / hr.
[0060]
The stirring tank type first polymerization vessel 17 is under the conditions of a reaction temperature of 235 ° C. and a reaction pressure of 98 mmHg. When the liquid volume of the molten polymer reaches 40 liters, the stirring tank type second polymerization vessel 17 is maintained so as to keep the liquid volume of 40 liters constant. An equal amount of molten polymer having a number average molecular weight of 850 was continuously supplied to the polymerization vessels 26a and 26b. The stirring tank type second polymerizers 26a and 26b are under conditions of a reaction temperature of 252 ° C. and a reaction pressure of 6 mmHg. When the liquid volume of the molten polymer reaches 20 liters, the horizontal biaxial shaft is used to keep the liquid volume of 20 liters constant. The molten polymer was continuously supplied to the stirring type polymerizers 35a and 35b. The hydroxyl group terminal ratio of the prepolymer monitored by sampling ports 52a and 52b and automatic viscometers 53a and 53b in the vicinity of the outlet of the stirred tank type second polymerization vessel was 43%, and the number average molecular weight was 2550. Since the target of the polycarbonate prepolymer produced in the stirred tank type second polymerization vessel 26a was 41% and 2600, Xbi = 0.0357 and Wbi = 0.17 were obtained from Equation 1 and Equation 2, and 0.017 kg / Hr of diphenyl carbonate was supplied in a molten state from the inlet 28a to the stirring vessel type second polymerization vessel 26a. On the other hand, since the target of the polycarbonate prepolymer produced in the stirring tank type second polymerization vessel 26b was 38% and 2600, Xbi = 0.0888 and Wbi = 0.046 were obtained from Equations 1 and 2. 0.046 kg / hr of diphenyl carbonate was supplied in a molten state from the inlet 28b to the stirred tank type second polymerization vessel 26b. Thereafter, the hydroxyl group terminal ratio and molecular weight were measured at intervals of 30 minutes for both the stirred tank type second polymerization devices 26a and 26b, and similarly calculated based on Equation 1, Equation 2, Equation 3, Equation 4, Equation 6 and Equation 7. Further, the hydroxyl group terminal ratio was continuously controlled by continuously supplying the amount of diphenyl carbonate or bisphenol A.
[0061]
In the horizontal biaxial stirring type polymerizers 35a and 35b, the reaction temperature is 265 ° C. and the reaction pressure is 0.6 mmHg. When the liquid volume of the molten polymer reaches 10 liters, the wire volume is maintained so as to keep the liquid volume 10 liters constant. Molten polymers having number average molecular weights of 6100 and 6000 were supplied to the perforated plate type polymerizers 42a and 42b, respectively. The perforated plate polymerizer 42a with wire is maintained at a reaction temperature of 265 ° C., and the reaction pressure is within a range of 0.4 ± 0.3 mmHg so as to obtain a target number average molecular weight of 10600 based on the measured value of the online viscometer 54a. Controlled by. When the liquid volume of the molten polymer in the lower part of the polymerization vessel reached 20 liters, the polycarbonate was extracted from the extraction port 51a through the transfer pipe 50a so as to maintain the liquid volume of 20 liters. The number average molecular weight of the obtained polycarbonate was 10600, and the hydroxyl group terminal ratio was 30 mol%. On the other hand, the perforated plate polymerizer with wire 42b is maintained at a reaction temperature of 272 ° C., and based on the measured value of the online viscometer 54b, the reaction pressure is 0.2 ± 0.2 mmHg so as to obtain the target number average molecular weight 12600. Controlled in the range. When the liquid volume of the molten polymer in the lower part of the polymerization vessel reached 20 liters, the polycarbonate was extracted from the extraction port 51b through the transfer pipe 50b so as to maintain the liquid volume of 20 liters. The number average molecular weight of the obtained polycarbonate was 12600, and the hydroxyl group terminal ratio was 17 mol%.
[0062]
The continuous operation was performed for 800 hours from the start. During this time, the molecular weight and hydroxyl group terminal ratio of the two types of polycarbonate measured every 3 hours were as follows: number average molecular weight 10600 ± 300 / hydroxyl terminal ratio 30 ± 1 mol%, number average molecular weight 12600 ± 400 / hydroxyl terminal ratio 17 Two types of polycarbonate having a molecular weight and a hydroxyl group terminal ratio in the range of ± 1 mol% could be stably and continuously produced without switching loss and quality fluctuation. The color tone of the obtained polycarbonate was also good.
[0063]
[Comparative Example 1]
In the process as shown in FIG. 2, two types of polycarbonates having different molecular weights and hydroxyl end ratios were produced by switching the production conditions. First, a polycarbonate having a number average molecular weight of 10600 / hydroxyl terminal ratio of 30 mol% was produced, and then the raw material initial charge molar ratio and the temperature of the final reactor were changed to obtain a polycarbonate having a number average molecular weight of 12600 / hydroxyl terminal ratio of 17 mol%. Manufactured. However, the apparatus used was the same as in FIG. 1, and only the series after 26a was used.
[0064]
The raw material dissolution and mixing tanks 3a and 3b are under the conditions of a reaction temperature of 180 ° C., a reaction pressure and a normal pressure, and a seal nitrogen (oxygen concentration of 0.5 ppm) gas flow rate of 1 liter / hr. 80 kg of bisphenol A powder and diphenyl carbonate powder (with a molar ratio of bisphenol A of 1.10) of 50 mmHg and vacuum nitrogen substitution 5 times were charged in the raw material dissolution mixing tank 3a and 7 mg of sodium hydroxide were melted for 2.5 hours. After mixing, the entire amount of the molten mixture was transferred from the transfer pipe 7a to the transfer storage tank 10. The melt mixture had undergone a transesterification reaction, and the reaction rate of bisphenol A was 70%. Similarly, the raw materials were similarly dissolved and mixed in the raw material dissolution and mixing tank 3b, and transferred to the storage tank 10 before the molten polymer in the storage tank 10 disappeared. The raw material dissolution and mixing tanks 3a and 3b repeated dissolution and mixing and transfer alternately. The storage tank 10 is maintained at normal pressure and 180 ° C. The molten polymer transferred to the storage tank 10 was continuously supplied to the stirring tank type first polymerization device 17 at 10 kg / hr.
[0065]
The stirring tank type first polymerization vessel 17 is under the conditions of a reaction temperature of 235 ° C. and a reaction pressure of 105 mmHg. When the liquid volume of the molten polymer reaches 20 liters, the stirring tank type second polymerization vessel 17 is maintained so as to keep the liquid volume of 20 liters constant. An equal amount of molten polymer having a number average molecular weight of 850 was continuously supplied to the polymerization vessel 26a. The stirring tank type second polymerization vessel 26a is under the conditions of a reaction temperature of 252 ° C. and a reaction pressure of 6 mmHg. When the liquid volume of the molten polymer reaches 20 liters, a horizontal biaxial stirring type is used to keep the liquid volume of 20 liters constant. The molten polymer was continuously supplied to the polymerization vessel 35a. The hydroxyl group terminal ratio of the prepolymer monitored by the sampling port 52a near the outlet of the stirred tank type second polymerization vessel and the automatic viscometer 53a was 40%, and the number average molecular weight was 2650.
[0066]
In the horizontal biaxial stirring type polymerizer 35a, the reaction temperature is 265 ° C., the reaction pressure is 0.6 mmHg, and when the liquid volume of the molten polymer reaches 10 liters, the porous volume with wire is maintained so that the liquid volume of 10 liters is kept constant. A molten polymer having a number average molecular weight of 5900 was supplied to the plate polymerization vessel 42a. The perforated plate polymerizer 42a with wire is maintained at a reaction temperature of 265 ° C., and the reaction pressure is within a range of 0.4 ± 0.3 mmHg so as to obtain a target number average molecular weight of 10600 based on the measured value of the online viscometer 54a. Controlled by. When the liquid volume of the molten polymer in the lower part of the polymerization vessel reached 20 liters, the polycarbonate was extracted from the extraction port 51a through the transfer pipe 50a so as to maintain the liquid volume of 20 liters. The number average molecular weight of the obtained polycarbonate was 9800, and the hydroxyl group terminal ratio was 23 mol%. Since the hydroxyl group terminal ratio was lower than the target polycarbonate, the raw material charge molar ratio 1.10 was changed to 1.08, and continuous operation was continued. Then, after changing the raw material charge molar ratio twice by trial and error, the raw material charge was finally set to 1.09 to obtain a polycarbonate having a number average molecular weight of 10600 and a hydroxyl group terminal ratio of 30%. During this time, about 100 kg of polycarbonate, which was different from the target, was produced and incurred. Further, continuous operation was continued under the same conditions, and the number average molecular weight and the hydroxyl terminal ratio were measured every 3 hours. The yield of the polycarbonate in the range of the number average molecular weight 10600 ± 300 / hydroxyl terminal ratio was 30 ± 1%. 75%.
[0067]
Next, in order to obtain a polycarbonate having a number average molecular weight of 12600 / hydroxyl terminal ratio of 17 mol%, the charged molar ratio is 1.08, the temperature of the perforated plate polymerizer 42a with wire is 272 ° C., and the pressure is 0.2 ± 0. The manufacturing was continued after changing to 2 mmHg. Thereafter, the charged molar ratio was changed three times to obtain the target polycarbonate, which was lost due to switching of about 250 kg of polycarbonate during this period. Thereafter, continuous operation for 800 hours was carried out, and the number average molecular weight and the hydroxyl group terminal ratio were measured every 3 hours. The number average molecular weight was 12600 ± 400 / the hydroxyl group terminal ratio was 17 ± 1 mol%. there were.
[0068]
[Example 2]
In the process as shown in FIG. 3, two types of polycarbonates having different molecular weights (number average molecular weight 8500 / hydroxyl terminal ratio 32 mol% and number average molecular weight 13000 / hydroxyl terminal ratio 25 mol%) were continuously produced simultaneously. However, the apparatuses 3a, 3b, 10, 17, 26a, 42a, and 42b used were the same as those in FIG. 1, and a perforated plate polymerizer 56 and a static mixer 67 of the same type as 42a were used.
[0069]
The raw material dissolution mixing tanks 3a and 3b are under the conditions of a reaction temperature of 182 ° C., a reaction pressure and a normal pressure, and a seal nitrogen (oxygen concentration of 0.5 ppm) gas flow rate of 1 liter / hr. 80 kg of bisphenol A powder and diphenyl carbonate powder (1.08 molar ratio of bisphenol A) 1.0 mol of sodium hydroxide was charged in the raw material dissolution and mixing tank 3a with vacuum nitrogen substitution at 50 mmHg 5 times and melted for 2.5 hours. After mixing, the entire amount of the molten mixture was transferred from the transfer pipe 7a to the transfer storage tank 10. The melt mixture had undergone a transesterification reaction, and the reaction rate of bisphenol A was 78%. Similarly, the raw materials were similarly dissolved and mixed in the raw material dissolution and mixing tank 3b, and transferred to the storage tank 10 before the molten polymer in the storage tank 10 disappeared. The raw material dissolution and mixing tanks 3a and 3b repeated dissolution and mixing and transfer alternately. The storage tank 10 is maintained at normal pressure and 182 ° C. The molten polymer transferred to the storage tank 10 was continuously supplied to the stirring tank type first polymerization device 17 at 20 kg / hr.
[0070]
The stirring tank type first polymerization vessel 17 is under the conditions of a reaction temperature of 235 ° C. and a reaction pressure of 105 mmHg. When the liquid volume of the molten polymer reaches 20 liters, the stirring tank type second polymerization vessel 17 is maintained so as to keep the liquid volume of 20 liters constant. An equal amount of a molten polymer having a number average molecular weight of 860 was continuously supplied to the polymerization vessel 26a. The stirring tank type second polymerization vessel 26a is under the conditions of a reaction temperature of 254 ° C. and a reaction pressure of 3 mmHg, and when the liquid volume of the molten polymer reaches 20 liters, a perforated plate type with wire so as to keep the liquid volume of 20 liters constant. A polymer melt having a number average molecular weight of 2330 was continuously supplied to the polymerization vessel 56. In the perforated plate polymerizer 56 with a wire, the reaction temperature is 260 ° C. and the reaction pressure is 0.8 mmHg. When the liquid volume of the molten polymer in the lower part of the polymerizer reaches 20 liters, the liquid volume is maintained to be 20 liters. It was supplied to a static mixer 67 kept at ° C. From the polycarbonate prepolymer and the in-line viscometer collected from the sampling port 68, the number average molecular weight of the molten polymer was 5500, and the hydroxyl group terminal ratio was 41 mol%. In order to obtain a polycarbonate having a target molecular weight and a hydroxyl group terminal ratio, it is necessary that the analysis value at the sampling port 68 is a number average molecular weight of 5500 and the hydroxyl group terminal ratio is 38 mol%. From Xci = 0.0605, Wci = 0.030 determined by use, 0.030 kg / hr of bisphenol A was supplied from the supply port 66 to the static mixer. Thereafter, the hydroxyl group terminal ratio and molecular weight were measured at 30-minute intervals, and the amount of diphenyl carbonate or bisphenol A calculated based on the formula 1, formula 2, formula 3, formula 4, formula 6 and formula 7 was continued. The hydroxyl terminal ratio was continuously controlled.
[0071]
The molten polymer discharged from the static mixer was supplied in an equal amount to the perforated plate polymerizers 42a and 42b with wires. The perforated plate polymerizer 42a with wire is maintained at a reaction temperature of 276 ° C., and the reaction pressure is 0.3 ± 0.3 mmHg so as to obtain a polycarbonate having a number average molecular weight of 13000 based on the measured value of the online viscometer 54a. Controlled in the range. When the liquid volume of the molten polymer in the lower part of the polymerization vessel reached 20 liters, the polycarbonate was extracted from the extraction port 51a through the transfer pipe 50a so as to maintain the liquid volume of 20 liters. The number average molecular weight of the obtained polycarbonate was 13000, and the hydroxyl group terminal ratio was 25 mol%. The perforated plate polymerizer 42b with wire is maintained at a reaction temperature of 260 ° C., and the reaction pressure is 0.6 ± 0.3 mmHg so as to obtain a polycarbonate having a number average molecular weight of 8500 based on the measured value of the online viscometer 54b. Controlled in the range. When the liquid volume of the molten polymer in the lower part of the polymerization vessel reached 20 liters, the polycarbonate was extracted from the extraction port 51b through the transfer pipe 50b so as to maintain the liquid volume of 20 liters. The number average molecular weight of the obtained polycarbonate was 8500, and the hydroxyl group terminal ratio was 32 mol%.
[0072]
The continuous operation was performed for 800 hours from the start. During this period, the molecular weight and hydroxyl group terminal ratio of the two types of polycarbonate measured every 3 hours were as follows: number average molecular weight 13000 ± 400 / hydroxyl terminal ratio 25 ± 1 mol%, number average molecular weight 8500 ± 300 / hydroxyl terminal ratio 32 Two types of polycarbonate having a molecular weight and a hydroxyl group terminal ratio in the range of ± 1 mol% could be stably and continuously produced without switching loss and quality fluctuation. The color tone of the obtained polycarbonate was also good.
[0073]
【The invention's effect】
In the method of the present invention, as shown in detail in the examples, the polycarbonate obtained can be stably produced with little molecular weight variation, and the switching loss in producing a plurality of polycarbonates having different molecular weights and hydroxyl end ratios. Thus, a plurality of polycarbonates having a constant quality can be stably and continuously produced at the same time.
[Brief description of the drawings]
FIG. 1 is a diagram of an apparatus for simultaneous and continuous production of 22 types of polycarbonates having different molecular weights and hydroxyl end ratios.
FIG. 2 is a diagram of an apparatus similar to FIG. 1 and using only the series 26a and later.
FIG. 3 is a view similar to FIG. 1 and further showing an apparatus using a polymerization vessel 56 and a mixer 67 of the same type as 42a.
[Explanation of symbols]
1a, 1b Raw material supply port
2a, 2b, 9, 19, 28, 36, 46 Vent port
3a, 3b Raw material dissolution and mixing tank
4a, 4b, 18, 27a, 27b, 39a, 39b Stirring shaft
5a, 5b, 11, 20, 29a, 29b, 47a, 47b, 57 Molten polymer
6a. 6b, 12, 21, 30a, 30b, 37a, 37b, 48a, 48b, 64 52a, 52b, 68 outlet
7a, 7b, 13, 15, 24, 33a, 33b, 40a, 40b, 50a, 50b, 65, 70
8, 16, 25a, 25b, 28a, 28b, 34a, 34b, 41a, 41b, 55, 66 Supply port
10 Storage tank
14, 23, 32a, 32b, 38a, 38b, 49a, 49b, 58 Transfer pump
17 Stirring tank type first polymerizer
22, 31a, 31b, 45a, 45b, 63 Gas supply port
26a, 26b Stirring tank type second polymerizer
35a, 35b Horizontal type biaxial stirring type polymerizer
42a, 42b, 56 Wire perforated plate polymerizer
43a, 43b, 61 perforated plate
44a, 44b, 62 wire
51a, 51b Product outlet
53a, 53b, 54a, 54b In-line viscometer
67 Static mixer

Claims (2)

In a method of continuously producing a polycarbonate having a number average molecular weight of 5000 to 30000 by polymerizing in a molten state from an aromatic dihydroxy compound and diaryl carbonate using at least two reactors, the reaction temperature is 150 to 290 ° C. Polycarbonate prepolymer production process A having a pressure of 10 mmHg to normal pressure, reaction temperature of 180 to 270 ° C., hydroxyl group terminal ratio control step B of the reaction pressure of 10 mmHg to normal pressure, reaction temperature of 230 to 300 ° C., reaction pressure of 100 mmHg It consists of the following post-polymerization process C, and when a single B process or a plurality of B processes provided in parallel is connected after a series of A processes, a plurality of B processes are provided after the B process. C processes provided in parallel are connected, and a plurality of B processes provided in parallel are connected. The C process provided in step C or more parallel singular after each step B is connected to the case of providing a degree, and a hydroxyl group-terminated ratio control process B is the molecular weight of the polycarbonate prepolymer and a hydroxyl terminated After measuring the ratio, an aromatic dihydroxy compound or diaryl carbonate is added to the liquid of the polycarbonate prepolymer based on the result, and the reaction is performed, so that the difference between the target hydroxyl group terminal ratio and the hydroxyl group terminal ratio of the polycarbonate prepolymer is ± Polycarbonates with different molecular weights without changing the charged molar ratio of the aromatic dihydroxy compound and diaryl carbonate used as the initial raw material by controlling the degree of pressure reduction in the C step. Is manufactured at the same time. A method for continuously producing carbonate. The aromatic dihydroxy compound or diaryl carbonate to be added is added in the amount Wci of the aromatic dihydroxy compound or the amount Wbi of the diaryl carbonate obtained from the following formula 1, formula 2, formula 3, and formula 4. A method for continuously producing a plurality of polycarbonates having different molecular weights according to item 1. (However, among the obtained Wbi and Wci, a component calculated as a positive value is added).

(Wherein, OH ₀ is the prepolymer target terminal hydroxyl group ratio (mol%), Mn ₀ is the target number average molecular weight of the prepolymer, OH _a is measured terminal hydroxyl group ratio of the prepolymer, Mn _a is measured in the prepolymer the number average molecular weight, Mn _b is the molecular weight of the diaryl carbonate to be added, the molecular weight of the aromatic dihydroxy compound Mn _c is to be added, X _bi is the mole fraction of diaryl carbonate to be added at the point i, X _ci is added at the time i Mole fraction of aromatic dihydroxy compound, W _ai is the weight of prepolymer passing through the reactor at time i (g / hr), Wbi is the weight of diaryl carbonate added at time i (g / hr), Wci is time This is the weight (g / hr) of the aromatic dihydroxy compound added at i, where a is a device coefficient determined by trial and error, 0.5 to 1 It is in the range of 0.)

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