JPH0337915B2 - - Google Patents
Info
- Publication number
- JPH0337915B2 JPH0337915B2 JP57144000A JP14400082A JPH0337915B2 JP H0337915 B2 JPH0337915 B2 JP H0337915B2 JP 57144000 A JP57144000 A JP 57144000A JP 14400082 A JP14400082 A JP 14400082A JP H0337915 B2 JPH0337915 B2 JP H0337915B2
- Authority
- JP
- Japan
- Prior art keywords
- reaction
- fermentation
- compressor
- liquid
- temperature
- Prior art date
- Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
- Expired - Lifetime
Links
- 238000006243 chemical reaction Methods 0.000 claims description 61
- 239000007788 liquid Substances 0.000 claims description 45
- 238000000855 fermentation Methods 0.000 claims description 35
- 230000004151 fermentation Effects 0.000 claims description 35
- 238000000034 method Methods 0.000 claims description 19
- LFQSCWFLJHTTHZ-UHFFFAOYSA-N Ethanol Chemical compound CCO LFQSCWFLJHTTHZ-UHFFFAOYSA-N 0.000 description 21
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 description 6
- 238000010438 heat treatment Methods 0.000 description 5
- 239000000047 product Substances 0.000 description 5
- 239000003507 refrigerant Substances 0.000 description 5
- 240000004808 Saccharomyces cerevisiae Species 0.000 description 4
- 238000009833 condensation Methods 0.000 description 4
- 230000005494 condensation Effects 0.000 description 4
- 230000006835 compression Effects 0.000 description 3
- 238000007906 compression Methods 0.000 description 3
- 238000010586 diagram Methods 0.000 description 3
- 238000004821 distillation Methods 0.000 description 3
- 238000011084 recovery Methods 0.000 description 3
- 230000035939 shock Effects 0.000 description 3
- 239000007795 chemical reaction product Substances 0.000 description 2
- 238000007796 conventional method Methods 0.000 description 2
- 238000000926 separation method Methods 0.000 description 2
- 230000002411 adverse Effects 0.000 description 1
- 238000013459 approach Methods 0.000 description 1
- 239000012295 chemical reaction liquid Substances 0.000 description 1
- 238000000354 decomposition reaction Methods 0.000 description 1
- 238000013461 design Methods 0.000 description 1
- 230000006866 deterioration Effects 0.000 description 1
- 230000000694 effects Effects 0.000 description 1
- 238000012423 maintenance Methods 0.000 description 1
- 238000006116 polymerization reaction Methods 0.000 description 1
- 239000002994 raw material Substances 0.000 description 1
- 238000010992 reflux Methods 0.000 description 1
- 230000005070 ripening Effects 0.000 description 1
- 238000009834 vaporization Methods 0.000 description 1
- 230000008016 vaporization Effects 0.000 description 1
Classifications
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02E—REDUCTION OF GREENHOUSE GAS [GHG] EMISSIONS, RELATED TO ENERGY GENERATION, TRANSMISSION OR DISTRIBUTION
- Y02E50/00—Technologies for the production of fuel of non-fossil origin
- Y02E50/10—Biofuels, e.g. bio-diesel
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/50—Improvements relating to the production of bulk chemicals
- Y02P20/59—Biological synthesis; Biological purification
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P70/00—Climate change mitigation technologies in the production process for final industrial or consumer products
- Y02P70/10—Greenhouse gas [GHG] capture, material saving, heat recovery or other energy efficient measures, e.g. motor control, characterised by manufacturing processes, e.g. for rolling metal or metal working
Landscapes
- Preparation Of Compounds By Using Micro-Organisms (AREA)
- Vaporization, Distillation, Condensation, Sublimation, And Cold Traps (AREA)
Description
この発明は、発酵反応のような発熱反応系の反
応熱を回収、有効利用するに当り、これを、反応
温度を制御しつつ最小所要エネルギーで行うよう
にした発酵反応熱回収方法に関する。
従来発熱反応系における反応熱の回収利用には
種々の試みがなされている。例えば、減圧発酵
槽中の反応温度の制御を、蒸発潜熱除去と、発生
蒸気の圧縮蒸気による加熱とで行う方法がある
（Vacuferm fermentation process；
Biotechnology and Bioengineering，Vol
XII pp 1749〜1751（1980））。しかし、この方法
は蒸発蒸気の圧縮に圧縮機を２段で用いるもので
あるが、圧縮蒸気を発酵槽の加熱源とし、生成物
であるエタノールをできるだけ多量に蒸発させる
ことを目的とするため、１段目の圧縮機の運転に
大量の電力（エネルギー）を消費し、２段目の圧
縮機は吸引力が低く、吐出側の冷却器による熱損
失が大きくなつているという欠点があつた。また
前述のように圧縮蒸気を発酵槽の加熱源としてい
るので加熱器壁の表層温度が高くなり発酵液に熱
シヨツクを与えることとなる。
また発酵槽から一部の発酵液を抜き出し、こ
れを熱交換器で加熱後、減圧下に保持したフラツ
シユ蒸発槽と精留塔に導き、生成物蒸気を蒸発さ
せ蒸気の圧縮蒸気を熱交換器の加熱源として利用
する一方、抜き出した発酵液は発酵槽に循環させ
る方法が提案されている（ATPAL法、Chem.
Age、Nov.21、p11（1980））。この方法は製品分
離塔（精留塔、ストリツピング塔）まで減圧下に
おくため定常運転が難しくなるという欠点があつ
た。またこの方法も圧縮機を２段にして用いるも
のであるが、製品分離塔の還流液に相当分の蒸気
も常時循環するので１段目の圧縮機による電力消
費が大量となり、アルコール収集タンクからの生
成物（エタノール）のロスが多くなるという欠点
がある。さらに前記の方法と同様発酵液に熱シ
ヨツクを与えるという問題も回避できなかつた。
発明者らはこのような従来法の欠点を克服し、
発酵反応熱の回収、利用をできるだけ少ない所要
エネルギーと簡単なシステム及び運転操作で達成
できる方法を開発するため種々検討を重ねてき
た。この結果、発生蒸気圧縮用の前、後段の圧縮
機の中間に間接熱交換器を設け減圧下の発酵反応
槽から反応熱に対応する量だけ発生させた蒸気を
前記圧縮機の１段目の圧縮機で圧縮し、この圧縮
蒸気を、減圧発酵反応槽からの発酵反応終了液と
前記間接熱交換器で熱交換させればその目的を満
足し得ることを見い出した。この発明はこの知見
に基づいてなされた。
すなわちこの発明は、発生蒸気圧縮用の前、後
段圧縮機の中間に間接熱交換器を設け、反応熱に
対応する量だけ減圧下の発酵反応器から取出した
蒸気を前段の圧縮機で圧縮して圧縮蒸気とし、こ
れを間接熱交換器において発酵反応終了液と熱交
換させ発酵反応終了液を昇温させることを特徴と
する発熱反応における反応熱回収方法を提供する
ものである。
以下図面に従がいこの発明を説明する。図面は
この発明方法の１実施態様のフローダイヤグラム
であり、図中１は発酵反応槽、２は発酵反応終了
液受槽、３は前段の蒸気圧縮機、４は後段の蒸気
圧縮機、５は間接熱交換器、６は気液分離機、７
はベントコンデンサー、８は回収反応熱で加熱後
の最終液受槽であり、９は発酵反応生成物分離の
ための蒸留塔である。
反応槽１には、ライン１０より発酵原料を供給
し、制御する温度で沸騰する圧力まで、蒸気圧縮
機３で反応槽１を減圧する。反応槽１から発生し
た蒸気をライン１１により蒸気圧縮機３に導入
し、所定の圧力、温度まで昇温、昇圧する。この
場合蒸気は反応熱に対応する量だけ反応槽１から
発生させる。反応槽１は単槽に限定されず、多槽
でもよい。
蒸気圧縮機３で圧縮された圧縮蒸気は、ライン
１２により間接熱交換器５に送られる。熱交換器
５には、反応槽１からライン１３により抜き出さ
れ、反応終了液受槽２に保持された反応終了液を
ポンプ１４を用いてライン１５から導入する。こ
の熱交換器５において反応終了液の加熱と反応終
了液を冷媒とする圧縮蒸気の凝縮を行う。熱交換
は向流で行うのが好ましい。この熱交換器５は中
間凝縮器として機能する。
なお、反応系（反応槽１）は減圧下にあるので
リーク・エアーは避けがたい。また非凝縮性ガス
が含まれる場合１段目の蒸気圧縮機３の吐出ガス
中に含まれる。したがつて熱交換器５で凝縮させ
た圧縮蒸気を気液分離機６に送り非凝縮性ガスを
分離する。
気液分離機６で分離された非凝縮性ガスは、ラ
イン１６から２段目の圧縮機４に送られ加圧され
たのちライン１７のベントコンデンサー７に送ら
れ、非凝縮性ガス中の蒸気をさらに凝縮回収後排
気される。圧縮機４としては真空ポンプが好まし
く用いられる。
最終液受槽８には間接熱交換器５で熱交換され
加熱された反応終了液がライン１８から送り込ま
れる。また、この受槽８には気液分離機６及びベ
ントコンデンサーの凝縮液もライン１９，２０よ
り送り込まれる。
最終液受槽８の昇温された反応終了液は、ポン
プ２１を有するライン２２により蒸留塔９に送ら
れ反応生成物が分離される。
この発明において、蒸気圧縮機３及び圧縮機４
の動力（エネルギー）を最小限に抑えるため、反
応槽１からの蒸気は反応熱に相当する量だけしか
発生させないことが必要である。
１段目の蒸気圧縮機３は、間接熱交換器５を用
いるので、反応終了液を冷媒とする凝縮温度に平
衡な圧力まで昇圧すればよく、それ以上余計に加
圧しても動力が余分にかかるだけで望ましくな
い。
この前段の蒸気圧縮機３の吐出圧力及び吐出温
度は、その圧縮がポリトロピツク圧縮となるので
以下の式で規定される。
T2＝T1（P2／P1）〓-1/〓×1/〓p ……(1)
T2＞Tm＋ΔT ……(2)
（P1＝反応系の圧力
P2＝吐出圧力
T1＝反応系の温度
T2＝吐出温度（〓）
Tm＝間接熱交換器の冷媒としての反応終了液の
温度（〓）
△Ｔ＝間接熱交換器５設計上のアプローチ温度
（〓）
γ＝断熱係数
ηp＝ポリトロピツク効率）
次にこの発明方法の操作条件をエタノール発酵
を例にとつて説明する。
反応系の温度は従来の酵母では25〜37℃であ
る。圧力はこれと平衡な圧力であり、24〜
60Torrである。もちろん高温で活性を示す酵母
が開発されればこれより高い圧力、温度でも適用
できる。
前段の蒸気圧縮機３の吐出圧力は通常80〜
500Torrであり、好ましくは100〜250Torrであ
る。吐出圧力が80Torr未満では平衡温度が40℃
以下となり、発酵温度が25〜37℃の発酵系では凝
縮不十分となる。一方500Torrを越えるとポンプ
効率の低下をきたす。さらには吐出温度が発火温
度（423℃）以上となる場合があり、操作上危険
である。
吐出温度は吐出圧力及び圧縮機の効率によつて
決まるが通常100〜380℃の範囲である。
反応終了液は間接熱交換器に25〜37℃で供給さ
れ、約47〜60℃に昇温される。前段の蒸気圧縮機
の吐出ガスは前記のように100〜380℃であるが、
反応終了液で40〜65℃程度に冷却凝縮される。
気液分離機は圧力90〜300Torr、温度40〜65℃
である。
後段の圧縮機の吐出圧力は、ベントコンデンサ
ーの圧損分だけ大気圧より高くする。
この発明方法は上記の構成を有し、その特徴な
いしは効果を列挙すれば次の通りである。
発酵反応系から発酵反応熱に相当する量しか
蒸気を発生させないので圧縮機の動力を最小限
に抑えることができる。
前、後段の圧縮機の中間に間接熱交換器を用
いるので、前段の蒸気圧縮機では発酵反応終了
液を冷媒とする熱交換での凝縮温度に平衡な圧
力まで昇圧すればよく、したがつて圧縮比を小
さくでき電力節減をはかることができる。
系を簡略化できるので操作、メインテナンス
が容易であり、もちろん投資コストも少なくて
すむ。
反応系を圧縮蒸気で直接加熱しないので熱シ
ヨツクにより反応液に悪影響を及ぼすような恐
れがない。
間接熱交換器として向流方式を用いれば凝縮
温度を冷媒（反応終了液）の入口温度近くまで
下げることができ、非凝縮性ガスに同伴する平
衡蒸気量をより低く抑えることができるので、
後段の圧縮機の容量を削減でき、電力をさらに
節減できる。
以上のようにこの発明の方法によれば反応熱の
回収及び利用が最小の所要エネルギーで達成でき
る。その１例を、従来法との比較で下記表に示
す。
The present invention relates to a fermentation reaction heat recovery method for recovering and effectively utilizing the reaction heat of an exothermic reaction system such as a fermentation reaction, by controlling the reaction temperature and using the minimum amount of energy required. Various attempts have been made to recover and utilize the reaction heat in conventional exothermic reaction systems. For example, there is a method of controlling the reaction temperature in a vacuum fermentation tank by removing the latent heat of vaporization and heating the generated steam using compressed steam (Vacuferm fermentation process).
Biotechnology and Bioengineering, Vol.
XII pp 1749-1751 (1980)). However, this method uses a two-stage compressor to compress the evaporated steam, but the purpose of this method is to use the compressed steam as a heating source for the fermenter and evaporate as much of the product ethanol as possible. The disadvantages were that a large amount of power (energy) was consumed to operate the first stage compressor, the suction power of the second stage compressor was low, and heat loss by the discharge side cooler was large. Furthermore, as mentioned above, since compressed steam is used as the heating source for the fermenter, the surface temperature of the heater wall increases, giving a thermal shock to the fermentation liquid. In addition, a part of the fermented liquid is extracted from the fermenter, heated in a heat exchanger, and then guided to a flash evaporator and a rectification column maintained under reduced pressure.The product vapor is evaporated and the compressed vapor is transferred to a heat exchanger A method has been proposed in which the extracted fermentation liquid is circulated in a fermentation tank while being used as a heating source (ATPAL method, Chem.
Age, Nov. 21, p11 (1980)). This method had the disadvantage that steady operation was difficult because the product separation towers (rectification tower, stripping tower) were kept under reduced pressure. In addition, this method also uses a two-stage compressor, but since steam equivalent to the reflux liquid from the product separation column is constantly circulated, the first-stage compressor consumes a large amount of power, and the alcohol collection tank The disadvantage is that there is a large loss of product (ethanol). Furthermore, as with the above-mentioned method, the problem of applying a heat shock to the fermentation liquid could not be avoided. The inventors overcame these drawbacks of the conventional method,
Various studies have been conducted to develop a method that can recover and utilize fermentation reaction heat with as little energy as possible and with a simple system and operation. As a result, an indirect heat exchanger is installed between the front and rear compressors for compressing the generated vapor, and the steam generated in the amount corresponding to the reaction heat from the fermentation reaction tank under reduced pressure is transferred to the first stage of the compressor. It has been found that the purpose can be achieved by compressing the steam using a compressor and exchanging heat with the fermentation reaction completed liquid from the vacuum fermentation reaction tank using the indirect heat exchanger. This invention was made based on this knowledge. That is, this invention provides an indirect heat exchanger between the front and rear compressors for compressing generated vapor, and compresses the steam taken out from the fermentation reactor under reduced pressure by the amount corresponding to the heat of reaction with the former compressor. The present invention provides a method for recovering reaction heat in an exothermic reaction, which is characterized in that the compressed steam is converted into compressed steam, and this is heat-exchanged with the fermentation reaction finished liquid in an indirect heat exchanger to raise the temperature of the fermentation reaction finished liquid. The present invention will be explained below according to the drawings. The drawing is a flow diagram of one embodiment of the method of this invention, in which 1 is a fermentation reaction tank, 2 is a fermentation reaction finished liquid receiving tank, 3 is a vapor compressor in the previous stage, 4 is a vapor compressor in the latter stage, and 5 is an indirect Heat exchanger, 6 is gas-liquid separator, 7
is a vent condenser, 8 is a final liquid receiving tank after heating with the recovered reaction heat, and 9 is a distillation column for separating fermentation reaction products. A fermentation raw material is supplied to the reaction tank 1 through a line 10, and the pressure in the reaction tank 1 is reduced by a vapor compressor 3 to a pressure at which it boils at a controlled temperature. Steam generated from the reaction tank 1 is introduced into the vapor compressor 3 through a line 11, and heated and pressurized to a predetermined pressure and temperature. In this case, steam is generated from the reaction vessel 1 in an amount corresponding to the heat of reaction. The reaction tank 1 is not limited to a single tank, but may be multiple tanks. Compressed steam compressed by the vapor compressor 3 is sent to the indirect heat exchanger 5 via a line 12. The reaction-finished liquid extracted from the reaction tank 1 through a line 13 and held in the reaction-finished liquid receiving tank 2 is introduced into the heat exchanger 5 through a line 15 using a pump 14 . In this heat exchanger 5, the reaction-completed liquid is heated and the compressed vapor is condensed using the reaction-completed liquid as a refrigerant. Preferably, the heat exchange takes place countercurrently. This heat exchanger 5 functions as an intermediate condenser. Note that since the reaction system (reaction tank 1) is under reduced pressure, leakage of air is unavoidable. In addition, if non-condensable gas is contained, it is contained in the discharge gas of the first stage vapor compressor 3. Therefore, the compressed vapor condensed in the heat exchanger 5 is sent to the gas-liquid separator 6 to separate non-condensable gas. The non-condensable gas separated by the gas-liquid separator 6 is sent from the line 16 to the second-stage compressor 4, where it is pressurized, and then sent to the vent condenser 7 in the line 17, where the vapor in the non-condensable gas is It is further condensed and collected and then exhausted. As the compressor 4, a vacuum pump is preferably used. The reaction completed liquid, which has been heat exchanged and heated by the indirect heat exchanger 5, is fed into the final liquid receiving tank 8 from a line 18. Further, condensed liquid from the gas-liquid separator 6 and the vent condenser is also fed into this receiving tank 8 through lines 19 and 20. The heated reaction-finished liquid in the final liquid receiving tank 8 is sent to the distillation column 9 through a line 22 having a pump 21, where reaction products are separated. In this invention, the vapor compressor 3 and the compressor 4
In order to minimize the power (energy) of the reaction vessel 1, it is necessary to generate steam from the reaction vessel 1 only in an amount corresponding to the heat of reaction. Since the first-stage vapor compressor 3 uses an indirect heat exchanger 5, it is only necessary to increase the pressure to a pressure that is balanced with the condensation temperature using the reaction-completed liquid as a refrigerant, and even if the pressure is increased more than that, there will be no need for extra power. This is not desirable. The discharge pressure and discharge temperature of the vapor compressor 3 in the previous stage are defined by the following equations since the compression is polytropic compression. T 2 = T 1 (P 2 /P 1 )〓 -1/ 〓 ×1/ 〓 p ……(1) T 2 ＞Tm＋ΔT ……(2) (P 1 = Reaction system pressure P 2 = Discharge pressure T 1 = Temperature of the reaction system T 2 = Discharge temperature (〓) Tm = Temperature of the reaction finished liquid as a refrigerant of the indirect heat exchanger (〓) △T = Design approach temperature of indirect heat exchanger 5 (〓) γ = Adiabatic coefficient ηp=polytropic efficiency) Next, the operating conditions of the method of this invention will be explained using ethanol fermentation as an example. The temperature of the reaction system is 25-37°C in conventional yeast. The pressure is in equilibrium with this, and is 24~
It is 60 Torr. Of course, if yeast that is active at high temperatures is developed, it can be applied at higher pressures and temperatures. The discharge pressure of the vapor compressor 3 in the previous stage is usually 80~
500 Torr, preferably 100 to 250 Torr. When the discharge pressure is less than 80Torr, the equilibrium temperature is 40℃
In a fermentation system with a fermentation temperature of 25 to 37°C, condensation is insufficient. On the other hand, if it exceeds 500 Torr, the pump efficiency will decrease. Furthermore, the discharge temperature may exceed the ignition temperature (423°C), which is dangerous for operation. The discharge temperature is determined by the discharge pressure and the efficiency of the compressor, but is usually in the range of 100 to 380°C. The reaction-finished liquid is supplied to an indirect heat exchanger at 25-37°C, and the temperature is raised to about 47-60°C. As mentioned above, the discharge gas from the vapor compressor in the previous stage is 100 to 380°C.
The reaction completed liquid is cooled and condensed to about 40-65℃. Gas-liquid separator has a pressure of 90 to 300 Torr and a temperature of 40 to 65℃.
It is. The discharge pressure of the downstream compressor is set higher than atmospheric pressure by the amount of pressure loss in the vent condenser. The method of this invention has the above-mentioned configuration, and its characteristics and effects are listed as follows. Since the fermentation reaction system generates steam only in an amount corresponding to the heat of the fermentation reaction, the power of the compressor can be kept to a minimum. Since an indirect heat exchanger is used between the front and rear compressors, the steam compressor at the front stage only needs to raise the pressure to a pressure that is in equilibrium with the condensation temperature during heat exchange using the fermentation reaction finished liquid as the refrigerant. It is possible to reduce the compression ratio and save power. Since the system can be simplified, operation and maintenance are easy, and of course investment costs can be reduced. Since the reaction system is not directly heated with compressed steam, there is no risk that the reaction liquid will be adversely affected by heat shock. If a countercurrent method is used as an indirect heat exchanger, the condensation temperature can be lowered to near the inlet temperature of the refrigerant (reaction completed liquid), and the amount of equilibrium vapor accompanying the non-condensable gas can be kept lower.
The capacity of the downstream compressor can be reduced, further saving power. As described above, according to the method of the present invention, the recovery and utilization of reaction heat can be achieved with the minimum required energy. One example is shown in the table below in comparison with the conventional method.
【表】
この発明の方法は発酵反応ばかりでなく、発熱
反応の化学反応で、平衡圧が大気圧であり、また
重合、分解、系の劣化を伴うため反応終了時まで
温度を上げられない系にも適用できる。
次にこの発明を実施例に基づきさらに詳細に説
明する。
実施例
図示のフローダイヤグラムに従がいエタノール
発酵を行つた。
反応槽１（内容積1200）に原料発酵液（組成
糖分＝15.3％、水84.7％）1090を仕込み、酒
母を添加後、空気を吹込み33℃、39Torrで発酵
を行わせる。酒母添加後48時間発酵を行わせて組
成 残糖分＝2.8％、水＝89.6％、エタノール＝
7.6％の熟成もろみ970を得、これを抜き出して
反応終了液受槽２に移した。
この発酵操作を繰り返したが、各発酵操作では
蒸気圧縮機３を作動させ反応槽１を上記減圧条件
に保持した。
この時反応槽１から発生しヘツダー（図示しな
い）に集められた蒸気はエタノール19.2％、
CO20.9％、水79.9％の組成を有し、発生速度平均
＝1.6Kg／hrであつた。このエタノール蒸気を蒸
気圧縮機３で圧縮して235℃、250Torrの蒸気と
した後、間接熱交換器５に導き、ここで、前記
の、反応終了液受槽２から抜き出され、ライン１
５から送られてきた、熟成もろみ40.3Kg／hr
（35.4℃）との熱交換が行われ、熟成もろみは57
℃まで加熱される。一方エタノール蒸気はこの結
果冷却されて凝縮し、50℃、200Torrの混相流と
なり、凝縮液は気液分離機６で分離された。凝縮
液は組成エタノール18.2％、水81.8％、温度50℃
であり、1.5Kg／hrで最終液受槽８に送られる。
気液分離機６で凝縮しなかつた蒸気は圧縮機４で
再圧縮後、ベントコンデンサー７で冷却されて、
エタノール62.3％、水37.7％の凝縮液を0.04Kg／
hrで生成し、最終液受槽８に送り込まれる。これ
らの凝縮液は、ライン１８から、この最終液受槽
８に送り込まれた前記の57℃に昇温した熟成もろ
みと混合されて、エタノール8.0％、水89.1％、
その他2.9％の組成で温度56.7℃の熟成もろみが
41.9Kg／hrで得られた。[Table] The method of this invention is applicable not only to fermentation reactions, but also to exothermic chemical reactions, where the equilibrium pressure is atmospheric pressure, and which involves polymerization, decomposition, and system deterioration, so the temperature cannot be raised until the end of the reaction. It can also be applied to Next, the present invention will be explained in more detail based on examples. EXAMPLE Ethanol fermentation was carried out according to the illustrated flow diagram. 1090 ml of raw fermentation liquid (composition: sugar content = 15.3%, water 84.7%) is charged into reaction tank 1 (inner volume 1200 ml), and after adding the yeast mash, air is blown in to carry out fermentation at 33°C and 39 Torr. After fermentation for 48 hours after addition of yeast mother, composition Residual sugar = 2.8%, Water = 89.6%, Ethanol =
A 7.6% aged mash 970 was obtained, which was extracted and transferred to the reaction finished liquid receiving tank 2. This fermentation operation was repeated, and in each fermentation operation, the vapor compressor 3 was operated to maintain the reaction tank 1 under the above-mentioned reduced pressure condition. At this time, the steam generated from reaction tank 1 and collected in the header (not shown) was 19.2% ethanol.
It had a composition of 0.9% CO 2 and 79.9% water, and the average generation rate was 1.6 Kg/hr. This ethanol vapor is compressed by a vapor compressor 3 to form a vapor of 235°C and 250 Torr, and then led to an indirect heat exchanger 5, where it is extracted from the above-mentioned reaction finished liquid receiving tank 2 and line 1
Aged moromi 40.3Kg/hr sent from 5
(35.4℃), and the ripening mash is 57℃.
heated to ℃. On the other hand, the ethanol vapor was cooled and condensed to form a multiphase flow at 50° C. and 200 Torr, and the condensate was separated by the gas-liquid separator 6. The composition of the condensate is 18.2% ethanol, 81.8% water, and the temperature is 50℃.
The liquid is sent to the final liquid receiving tank 8 at a rate of 1.5 kg/hr.
The vapor that was not condensed in the gas-liquid separator 6 is recompressed in the compressor 4, and then cooled in the vent condenser 7.
0.04Kg/condensate of 62.3% ethanol and 37.7% water
hr and sent to the final liquid receiving tank 8. These condensed liquids are mixed with the above-mentioned aged mash that has been heated to 57°C and sent from line 18 to this final liquid receiving tank 8, and are mixed with ethanol 8.0%, water 89.1%,
Aged mash with a composition of 2.9% and a temperature of 56.7℃
Obtained at 41.9Kg/hr.
図面はこの発明の反応熱回収方法の１例を示す
フローダイヤグラムである。
符号の説明、１……反応槽、２……反応終了液
受槽、３……蒸気圧縮機、４……圧縮機、５……
間接熱交換器、６……気液分離機、７……ベント
コンデンサー、８……最終液受槽、９……蒸留
塔。
The drawing is a flow diagram showing one example of the reaction heat recovery method of the present invention. Explanation of symbols, 1... Reaction tank, 2... Reaction finished liquid receiving tank, 3... Vapor compressor, 4... Compressor, 5...
Indirect heat exchanger, 6... Gas-liquid separator, 7... Vent condenser, 8... Final liquid receiver, 9... Distillation column.
Claims (1)
接熱交換器を設け、反応熱に対応する量だけ減圧
下の発酵反応器から取り出した蒸気を前段の圧縮
機で圧縮して圧縮蒸気とし、これを間接熱交換器
において発酵反応終了液と熱交換させ発酵反応終
了液を昇温させることを特徴とする発熱反応にお
ける反応熱回収方法。1. An indirect heat exchanger is installed between the front and rear compressors for compressing the generated vapor, and the steam taken out from the fermentation reactor under reduced pressure in an amount corresponding to the reaction heat is compressed by the front compressor to become compressed steam. A method for recovering reaction heat in an exothermic reaction, characterized by exchanging heat with the fermentation reaction finished liquid in an indirect heat exchanger to raise the temperature of the fermentation reaction finished liquid.
Priority Applications (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| JP57144000A JPS5934889A (en) | 1982-08-21 | 1982-08-21 | Recovery of reaction heat |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| JP57144000A JPS5934889A (en) | 1982-08-21 | 1982-08-21 | Recovery of reaction heat |
Publications (2)
| Publication Number | Publication Date |
|---|---|
| JPS5934889A JPS5934889A (en) | 1984-02-25 |
| JPH0337915B2 true JPH0337915B2 (en) | 1991-06-07 |
Family
ID=15351979
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| JP57144000A Granted JPS5934889A (en) | 1982-08-21 | 1982-08-21 | Recovery of reaction heat |
Country Status (1)
| Country | Link |
|---|---|
| JP (1) | JPS5934889A (en) |
-
1982
- 1982-08-21 JP JP57144000A patent/JPS5934889A/en active Granted
Also Published As
| Publication number | Publication date |
|---|---|
| JPS5934889A (en) | 1984-02-25 |
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