CA1243471A

CA1243471A - Process for carrying out reactions and mass transfer processes in heterogeneous fluid systems

Info

Publication number: CA1243471A
Application number: CA000475981A
Authority: CA
Inventors: Waldemar Reule; Georg Schreiber
Original assignee: V L T Gesellschaft fur Verfahrenstechnische Entwicklung Mbh
Current assignee: V L T Gesellschaft fur Verfahrenstechnische Entwicklung Mbh
Priority date: 1984-03-08
Filing date: 1985-03-07
Publication date: 1988-10-25
Also published as: DE3408464A1; EP0154334A3; EP0154334B1; EP0154334A2; ATE49353T1; BR8501059A; DE3575264D1; JPS6121727A; ZA851712B

Abstract

ABSTRACT OF THE DISCLOSURE:

A process for carrying out reactions and mass transfer processes in heterogeneous fluid systems, particu-larly solid-liquid systems and solid-liquid-gas systems.
The process comprises the utilization of a reactor with agitator vessel characteristic as first stage and of a reactor having a smaller volume with tube characteristic as second stage. The two reactors are in communication with each other on the basis of a communicating hydraulic system in the region of their upper and lower ends by pipe-lines, or the like.

Description

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The present invention relates to a process for ~he continuous performance of reactions and mass transfer processes in heterogeneous fluid systems, particularly solid-liquid systems and solid-liquid-gas systems.
It is known that in the case of chemical and biochemical reactions and in the case of mass transfer processes, such as adsorption, ion exchange, absorption, extraction, crystallization, rectification, filtration, etc., the residence time distribution exerts a substantial influence on the apparatus and plant dimensions, on turnover, yield, and selectivity as well as on the consumption of operating material and energy. It is also known that in a reaction apparatus without axial back mixing, in a socalIed ideal tube reactor, which is characterized by plug flow, in the case of simple isothermal reactions continuously a higher yield can be reached than in a thorougly mixed reactor of the same volume, for example an agitator vessel reactor, whereby instead of a tube reactor often with the same success also a vessel cascade may be utilized, consisting of a plurality of stirring apparatuses connected in series.
In certain cases it may also be necessary to combine an apparatus with agitator vessel chracteristic as the first stage with an apparatus with tube characteristic as the second. E. g., that is the case, if a reaction takes place very rapidly at the begining and results in extreme heat development. In order to eliminate the heat of reaction, it is necessary to produce, by proper mixing, a high heat transmission on the cooling surfaces. If it is a reaction, where ~

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the reaction rate i5 proportional to the concentrat~on of one or a plurality of reactantds, then, for a turnover as complete as possible, however, in most cases a very large apparatus volume of the thoroughly mixed reactor is necessary.
In a known manner this disadvantage may be avoided that the reaction is carried out just partly in the stirring appa-ratus and the reaction solution is supplied continuously from the stirring apparatus as first stage to an apparatus ' with tube characteristic as second stage, in which the reaction then takes place to the desired final conversion on the basis of a smaller apparatus volume. This distribution particularly is of advantage in the case of turnovers of more than 90 percent, due to the fact that in such a case the differences in the volume between reactors with agitator vessel character-istic and reactors with tube characteristic are significant.
There is the problem in the case of many continuousreactions and mass transfer processes in heterogeneous systems to separate one component, e.g. a reaction component or an auxiliary substance, such as suspended catalyst particles, from the reaction solution, after the reaction stage and/or transfer stage, respectively, and to lead back again into the process. Often extensive expenditure is required for this, particularly if such separation is to be carried out at pressure and at high temperatures, in order to avoid high loss of energy. Then a returning of solid material, e.g.
of catalyst particles, can be carried out only by additional expensive conveying equipment.
It is the object of the present invention to elimi-nate this disadvantage.
In meeting this and other objects, the invention provides a process for carrying out reactions and mass transfer processes in heterogeneous fluid systems, particularly solid-liquid systems and solid-liquid-gas systems, characterized by j~:
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' :: `:' 1~3 the utilization of a reactor with agitator vessel characteristic as first stage and of a reactor having a smaller volume with tube characteristic as second stage, the two reactors being in communication with each othcr on the basis of a communicating hydraulic system in the region of their upper and lower ends by pipelines, or the li~e.

By the fact that the two stages are in communication with each other by pipelines, they are a sing]e hydrau-lic communicating system. In the first stage in a known manner, for the purpose of thorough mixing, energy is introduced, for example hydromechanically by means of a stirrer, hydrostatically by the appli-cation of gas, or hydrodynamically by a driving liquidjet; this shows itself by kinetic energy of the fluid in the first reactor and therein leads via the moveme~t of the fluid in the form of flows to a pressure field, changed in regard to the s~tic condition. Thereby partial spaces with higher static pressure and partial spaces with lower static pressure are formed. In the first reactor, consequently in the first stage, thereby fluid streams from the spaces of higher pressure to the spaces of lower pressure, and owing to the momentum transfer by the stirring means, i. e., for example, stirrer, gas blowing, or power jet, on the basis of pressure build-up back to the spaces of higher pressure.

If a space of higher pressure is connected, for example, by a pipeline outside the first reactor, to a space of lower pressure, then fluid streams in this pressure gradient through the connecting line, con-sequently in parallel with the return streams in the first reactor. This is utilized by the present invention in that same incorporates into the connection between the spaces of higher pressure and lower pressure in the .

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/,`h, first reactor a second reactor, which, for the p~rpose of the reaction and/or the mass transfer, respectively, has tube characteristic. By this it is reached that a partial flow of the fluid multiphase system, which has taken part in reactions and/or mass transfer, respectively, in the first reactor under the conditions of more or less complete mixing, flows solely due to the momentum effect in the first reactor, that means without addi-tional supply of further external energy, under con-ditions of plug flow through the second stage and,under the influence of the same forces flows back again to the first stage, the intended far-reaching conversion and/or mass transfer, respectively, taking place therein.
The tendency of the partial flow to stream back without further supply of energy, consequently e. g.
without a pump, solely by the pressure difference be-tween the spaces of higher pressure and of lower pres-sure through the second stage and back to the firstlS lnCreaSed, stage,lif in a heterogeneous multiphase system before the position, where the partial flow of the fluid leaves from the first stage via the upper connection into the second stage/a phase separation occurs in such a manner that the fluid (e. g. a gas-liquid disper-sion or a suspension) has in this partial flow, con-sequently also in the second stage, a higher mean density than the polyphase system in the first stage and/or in the return flows, respectively, from spaces of higher pressure to spaces of lower pressure within the first stage.

For instance, such a case is siven, if the first stage is a bubble column reactor or a mammoth loop reactor, wherein a gas supplies the mixing energy and wherein the gas phase providing the kinetic energy (disperse phase) , prior to the entering of the fluid lZ~3~7~

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into the upper connectior-, leaves the multiphase system at the upper end of the ~irst stage. In this case the partial flow through the second stage has a higher mean density than the mixture in the bubble column reactor or in the portion of the circulation flow streaming downwards in the loop reactor.

The inlet, in the following called inset flow, may be supplied to the first stage or also to the second stage, whereas the outlet, in the following called product flow, is separated from the second stage, actually at a position, where the desired conversion is reached.

Further advantageous configurations and develop-ments of the invention are descrlbed in the sub-claims.

The invention is illustrated on the drawings in the form of embodiments and will be described on the basis of such embodiments in the following.

Fig. 1 shows, in schematical representation, one embodiment of the apparatus, required for carrying out the process of the present invention, that means, of the two-stage equipment, Fig. 2 in a simplified schematical view a variant, wherein the inset flow is delivered to the second stage, Fig. 3 a further variant, wherein the product flow is led out in a particular separating device, Fig. 4 an embodiment with internal circulation in the first stage, provided as a mammoth ~ loop reactor, ,, , .. :.

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i'"-3~ Fig. 5 an embodiment with outer circulation in the first stage, provided as a mammoth loop reactor, Fig. 6 the integration of the second stage into the frist stage, Figs. 7 two embodiments with t.he second stage, pro-and 8 vided as whirl cell tower, Fig. 9 an embodiment with two reactors as second stage, Fig 10 an embodiment with tandem connection of two first and second stages, Fig. 11 a schematic view of a plant for the operation of the process of the present invention by an:
example of an anaerobic processing of highly burdened effluents, Fig. 12 a further schematic view of a plant for the operation of the process of the present invention, whilst Fig~ 13 shows the application of the present invention in the form of a block diagram by example of the processing of an acidic effluent.

According to Fig. 1 the reactors 1 of the first stage and 2 of the second stage are connected with one another in the region of their upper ends by means of pipeline 3 and in the region of their lower ends by means of pipeline 4. As is indicated by arrows 5, the reactor 1 is a thoroughly mixed apparatus with agitator vessel characteristic, whilst reactor 2 is a reactor ,.

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The inset flow is introduced at 6, and the mixing energy, for example gas, at 7. In such caseS where gas is formed during the reaction, of course, such gas may be utilized, according to the known principle of the mammoth pump, as impulse for the mixing operation.
The product flow is taken away at 8, and the reaction gas is removed at 9. The level of the fluid mixture is designated by 10.

The fluid mixture introduced at 6 is caused to mix by the mixing energy supplied at 7, as symbolically indicated by the arrows 5. The two reactor stages 1 and 2 are arranged and designed in such a manner that a partial flow of the mixed fIuid mixture streams from the reactor 1 via the line 3 into the reactor 2 and streams through same by necessity in the form of a plug flow owing to its configuration with tube characteristic. Whilst at the position 8, where the desixed conversion is reached, the product flow is removed, the recirculated components flow back via the line 4 into the reactor 1 and take part in the further procedure of the process.

As is evident by Fig. Z, wherein same parts are again designated with the same reference numerals, the separation in the second stage may be carried out for example by the exchange of the fluid phase, con-taining the product for another fluid phase, contain-ing the inset. The mixing in the first stage, i. e.
in the reactor 1, is carried out in this case by means of the agitator 12 by the mixing drive 11.

As shown by Fig. 3, the product removal may be carried out also in a separate dividing device 13, arranged in the return line 4, that means outside the -` ~LZ~3~

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reactor 2, for example, by sedimentation, filtration, or the like.

The reactor 1, provided as a mammoth loop reactor (in the following designated with MSR) of the first stage is provided, according to Fig. 4, as a cylin-drical reactor having a coaxially installed guide tube 14 for the inner circulation and, according to Fig. 5, as a bubble column with outer return lead 15 for the outer circulation according to arrows 5.
Thereby two tube regions or partial spaces each are formed, through which flow -takes place in the form of a closed loop. The supply of gas, as a ruIe into the larger space of the two partial spaces, effects the circulation of the 1iq~id ~r o~ the suspensi~n, re-~pectively, the loop flow. The gas bubbles, disper~ed in the continuous phasè, bubble up in the liquid (suspension) due to their buoyancy ~potential energy).
Thereby impulse in the form of frictional and sweep-ing forces is transferred to the surrounding fluid,which, thus, flows in the direction of the gas. On the surface of the mixture gas bubbles emerge from the fiuid, however, the fluid flows downwards with a lower content of gas bubbles (or free of bubbles, ac-cording to the possibility of the degassing for thebubbles) back in the other partial space.

These processes of the impulse transfer between gas bubbles and fluid and the emerging of a gas flow from the fluid in the stationary condition are characterized by a difference of the gas content be-tween the two partial spaces, which also exists, if in the case of a reactor, wherein gas results by the reaction, the external gas supply is stopped. In the partial space with the larger volume the flow directed upwards is maintained due to the fact that therein ;

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~ more gas is produced than in the smaller partial space, r where the flow is directed downwards. The resultlng gas emerges on the surface of the mixture.

The gas content difference between the two ~artial spaces is equivalent to a difference of the static pressures between the two spaces (measured on the same level). Consequently, thereby a loop flow according to the pressure field results.
According to the present invention this pressure difference is utilized in that a partial flow of the liquid (suspension) is conveyed from the first-stage through the second stage of the device without further supply of ener~y.

In the case of the embodiment of Fig. 6 the reactor 2 of the second stage is integrated in the reactor 1 of the first stage so that in this case the lines 3 and 4 are not required. It goes without saying that the product removal is to be carried out also in this case in the region of the integrated reactor 2 at a suitable position.

According to Fig. 7 the reactor 1 is provided again as an MSR ('Mammutschlaufenreaktor'/mammoth colu~n reactor) with inner circulation and the reactor

2 is provided as a whirl c,ell tower /Wirbelzellen-kolonne/ (in the following called WZK) with the in-clined bottoms 16. In this case the reactor 1 is pro-vided as a reactor in connection with a reaction with the formation of gas, i. e., in this case it is done without the input of external mixing energy. In order to promote the flow in the reactor 1 the reactor bottom 17 is provided as a deviating bottom; primarily this is i~portant for suspensions. In addition, the line 18 ~.

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is provided at the upper end of the WZK fox the re-moval of reaction gas. According to Fig. 8, in a variant, in addition, line 19 with pump 37 is pro-vided, through which a partial flow i~ taken from the MSR at a suitable location and is supplied to the WZK, at the bottom, in order to exchange the continuous phase, and/or to return suspended particles into the MSR 1, respectively. On the presupposition that a dif-ference of pressure exists at the input position of the line 19 into the reactor 2 between line 19 and reactor 2, it is also possible to do without the use of the pump 37.

The embodiment of Fig. 9 shows, in connection with a reactox 1 of the first stage, two reactoxs 2 and 2' of the second stage. The inset flow 6 is supplied to the one reactox 2 of the second stage; this has the advantage that the inset flow conveys the fluid mixture parts, present in the lower portion of this reactor, further into the reactor 1. Of course, how-ever, it is also possible to carry out the inset 6 in this case directly into the reactor 1.

According to Fig. 10, two reactors 1 and 1' of the first stage are provided, one reactor 2 and 2' of the second stage each being associated therewith. In this connection it is arranged that the partial flow of the reactor 2 streams into the reactor 1' and the partial flow 2' stxeams into the reactor 1. Deviating from the illustration as shown, it may also be arranged that the inset flow is supplied just to the reactor 1 and the product flow is removed just from the reactor 2'.

Fig. 11 shows for the example of an anaerobic biological purification of waste water as first stage the NSR 1 with guide tube 14, and as second stage the ;~
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WZK 2 with inclined bottoms 16, both being connectedto one another in the region o~ their upper ends by means of line 3 and in the region of their lower ends by means o line 4. In addition, the MSR 1 and the WZK 2 are in communication with each othcr by means of the partial flow exchange l.ine 19. The waste water supply 6 into the MSR 1 is carried out by means of the waste water pump 22 via the waste water preheater 23, which simultaneously serves as pure water cooler, as ;
well as via the steamheated waste water heat exchanger 24. From the nutrient solution container 20, via the dosing pump 21, simultaneously the nutrient solution is supplied into the MSR 1 at 6. The supply of the mixing energy, for example of a gas such as nitrogen, is effected at 7. However, instead of that, it could also be provided that biogas from the circulation, after condensing, is used for the impulse. The removal of the product 8 from the WZK 2 is carried out via the pure water cooler 23, serving simultaneously as waste water preheater, and the pure water solid mat:ter sepa-rator 27, via the line 31 the pure water being obtained and via the line 32 the surplus sludge being:obtained.
The reaction gas is removed from the MSR 1 via the line 9 and from the WZK 2 via the line 18 and is supplied to the gas cooler 25 as well as to the condensate separator 26, wnereafter the biogas is removed via the line 29 and the condensate is returned into the MSR 1 via the line 28.

According to conditions prevailing in practical operation, there may also be cases, where the multi-phase system in the second stage has a lower mean density than the mixture in the first stage. The dif-ference of the mean densities of the multiphase system in the two stages even may be so significant that the difference caused thereby of the hydrostatic pressures ..... ...

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between the two stayes cannot be compensated any more by the pressure difference, induced in the first stage by impulse transfer between the spaces of higl pressure and low pressure.

For instance, this case may occur, if in such a two-stage reactor system a suspension shall circulate through the two stages. Particularly in the starting phase of the reactor system, if possibly the particles are filled merely into the first stage, however, also during the stationary operation, caused by certain lo-cal instabilities of the suspension flow, the mean density of the multiphase system may be or may become smaller in the second stage than in the first stage.
In such a case the suspension partial flow from the first stage through the second stage and back into the first stage would not be initiated or would come to a standstill.

However, it turned out in such a case (Fig. 12), where, in addition, to the liquid, continuous phase a heavier, disperse phase exists, consequently in the case of a suspension of particles, which possess a higher density than the liquid, that also in the space of the second stage, which is situated above the inlet position of the partial flow 3 into the second stage, a separation of the suspension occurs.

If the two stages, as shown in Figures 4, 5, 7, 8, and 11, are connected to each other via further lines 9 and 18, which end into the first stage above the mixture level 10, then also in the second stage, or in the line 18, respectively, a mixture level is formed, the level of which, however r is above the mixture level 10. This position is caused in that in the space of the second stage above the inlet position :

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-~ ~ of the partial flow 3, duc to the sedimentation of the heavier disperse phase, from the inlet position downwards merely the continuous liquid phase exists, whilst in the space of the first stage above the position, where the partial flow from the first stage leaves, a mi~ture consisting of continuous phase and heavier disperse phase exists. Due to the fact that the corresponding spaces of the two stages hydraulically communicate with each other, in the space with the smaller fluid density the liquid level must be higher than in the space with the higher density. This dif-ference in level is also present, if the mean density of the suspension in the second stage is larger than in the first stage. However, same is particularly large, if the second stage is filled only with the lighter liquid.

If the space above the mixture level 10 in the first stage is connected via a line 30 to the space of the second stage, which is situated above the zone of separation, consequently above the inlet position of the partial flow 3 into the second stage, then the lighter liquid flows through this line 30, under the action of the hydrostatic pressure of the liquid column, which is situated abo~e the level 10 in the first stage, back to the first stage. By this the hydrostatical pressure at the inlet position of the partial flow 3 into the second stage is reduced; thereby more sus-pension flows from the first stage via the line 3 into the second stage. The partial flow 30 of the liquid is maintained by the pressure difference, which is main-tained between the initial and end points of the line

3 owing to the density difference of the suspension column or liquid column, respectively. By the separation of the liquid-flow 30 from the suspension-flow 3 at the upper end of the second stage the suspension is .

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Z~3~t71 ~t concentrated (thickened) at this location so that the concentration of -th~ suspension in regard to the heavy phase in the second ~'-age become~ larger and larger when starting, and in the stationary operation is always larger than in the first stage. Thus, also the mean density in the second stage bccomes or is always larger than in the first stage and, thus, a reliable operation regarding the circulating partial flow through the second stage is secured. By throttling of the liquid flow 30, it is possible to regulate the suspension flow 3 and, thus, the concentration of the heavier disperse phase in the second stage.
, In Fig. 12 also the case is shown, where the bio-gas is condensed by means of the condensing device 39 and is utilized for the propulsion of the circulation in the MSR.

In the diagram for the processing of acid-laden process water according to Fi~. 13 the waste water containing acetic acid in this case is supplied through the line 35 initially the ion exchanger 33, from where a part of the waste water with CO2 is removed as puri-fied water via the line 34, and another portion of the waste water with CHCOONa flows to ~he MSR 1, which again is in communication with the W~K 2 by means of the lines 3 and 4. Whilst on the upper end of the MSR 1 and of the WZK 2, via the lines 9 and/or 18, respectively, as product CH4 ~ CO2 are removed, 30 waste water with NaHCO3 is taken from the WZK 2 at -~
a suitable location at 8 and is supplied to the ion exchanger 33 so that sodium ions are removed from the waste water and are returned into the reaction.

According to the process of the present invention the following systems may be operated with special advantage~

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'J' ~ ~ Solid-liquid systems with, at least, one solid phase, dispersed in, at least, one liquid, con-tinuous phase, as well as - gas-liquid-solid systems wi-th, at least, one gas phase and, at least, one solid phase, dispersed in, at least, one liquid, continuous phase.

In this connection, the following are taken into consideration, for the first stage:-- agitator vessels as-loop reactors with propeller agitator (propeller loop reactor/Propeller-Schlaufen-reaktor/: PSR) - loop reactors with propulsion of the circulation by mammoth pump ~"Airlift" mammoth loop reactor/
"Airlift" Mammutschlaufenreaktor/- MSR) - loop reactors with propulsion of the circulation by liquid jet (jet loop reactor/Strahl-Schlaufen-Reaktor/: SSR) - bubble columns - agitator vessels, and for the second stage:

- whirl cell towers or similar multistage cascade columns with self-fluidizing cells (stages) - columns with installations consisting of static mixers - sieve tray columns 3~71 .~
~ - packed columns ('~iillkorper-Kolonnen') .. ~ . .
- packed towers.

The advantageous applicability of the present invention will be described in the following in more detail on the basis of two new processes:-Process 1 Anaerobic processing of organically highly burdenedprocess waters, containing particularly low fatty acids and alcohols.

Numerous process water flows in the chemical and petrochemical industries, the food-processlng industry, - the ceIlulose and paper industries contain organic substances in such a high concentration that their processing or 'Entsorgung' (removal) is a tremendous problem. Often they are burnt at high costs, or are considerably diluted by other waste waters, and are then aerobically purified in biological waste-water purification plants, also wlth conslderable costs.

Due to the fact that the organic substances of the process waters, in principIe, represent re-usable raw materials and/or energy raw materials,it is ap-propriate and practical to.apply processinq methods, wherein these substances are reclaimed directly or are converted into a utilizable product, and, simul~
taneously, the water is purified to such an extent that the water can be removed or can be reused as process water.

The physical-chemical treatment of such process waters for the direct recuperation of the organic -' ' , .
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3~71 content substances is normally uneconomical due to the low concentration in regard to such processes.
Therefore, the application of fermentation processes for the conversion of the organic content substances into utiliæable products is an appropriate way for the processing of the process waters, if the attempt is successful to control and govern in process-tech-nical respect the, for biotechnical processes, high concentrations of the content substances (components).
This object is solved by the process of the present invention.
.
It is known that from anaerobic micro-organisms (normally bacteria) organic water components (dissolved and suspended organic substances, such as, e. g., fatty acids, alcohols, aldehydes, carbohydrates, proteins) may be utilized as carbon sources and energy sources and may be decomposed into "biogas"
(methane and carbon dioxidej. "Methanogeneous sub-.
strates" (acetic acid, formic acid, methanol, hydrogentand carbon dioxide)are converted by methane bacteria directly into methane (CI~4) and carbon dioxide (CO2).
All other substances must be converted initially by other groups of bacteria Ifermentative, hydrolyzing, and acetogenic bacteria) into methanogenic substrates.

For the technical execution of an anaerobic process one-stage and two-stage processes are known.
In the one-stage process all the decomposition reactions with the different groups of micro-organisms take place in a reactor, in the case of the two-stage process it is hydrolyzed and fermented (acidulated) in the first stage, and in the second stage the higher fatty acids from the first stage are converted by acetogenic bacteria into methanogenic substrates, which are con-verted then, together with the methanogenic substrates (mainly acetic acid, hydrogen, and CO2), also formed .

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`' already in the first stage, in a so-called "bio-energetic symbiosis" with the acetogenic bacteria by methane bacteria into biogas.

Accordin~ to exl~erience , it is necessary, in the second stage, as also in a continu~us process, ~herein methanogenic substrates are converted directly in one stage into biogas (e. g. acetic acid-laden was~e water, such as, for instance, water vapour condensate from the evaporation of sulfite waste liquors or reaction water from the production of dimethyl-terephthalate (DMT) and terephthalic acid (TPA)~, to keep the pH value in the reactor within narrow limits constant, as close as possible to an optimum value of ca. 6.8. This condition is of great importance, if one takes into consideration that in the outlet of the first stage of a biogas plant the pH value, in the case of an acetic acid concentration of approximately 1 moI/l, may be around 4, and that a process waste water of the DMT production may comprise a pH value of 2 - 4. These values considerably impede the optimum operation of a methane stage.

The adjustment of the optimum pH value for the methanization can be reached, if the conversion is carried out in a reactor of the type of an agitator vessel reactor, wherein "ideal mixing" as far as pos- ;
sible is secured. Thus, in the case of a determined conversion of the acids into C~14 and CO2 in the con-tinuous operation via the acid-base equilibria (CO2 or "carbonic acid" has a lower acidity than acetic acid), a constant pH value in the entire reactor volume may be adjusted. In this connection, the con-version is adjustable by the mean residence period of the throughput. In order to reach the optimum pH value - in this manner, possibly just a partial neutralization of the acids is necessary. ~ ;~

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,~, In a direc t connection with this, there is also ~~~ the possibility to supply an inlet flow having a relatively high content of organic substrate to the reactor. Sam~ is mixed with the content of the reactor almost immediately. Therefore, the mixture concen-tration, in the case of high conversion, is considerably lower thain the inlet concentration and locally uniform in the reactor. It depends on the hydraulic residence period and the converting rate of the micro-organisms (reaction rate). As for almost all the microbial re-actions, in the case of the anaerobic methane formation, the velocity of the substrate decomposition is coupled with the growth rate of the methane bacteria. However, same i5 very small as compared with aerobic cells.That means, in a continuously operated methane reactor an organically highly burdened inlet can be treated only with very long residence time with reasonable conversion, because, if the residence time is too short, normally the addition of cells in the reactor (with freely sus-pended anaerobesJ does not keep step with the outputby the throughput. In this case, the stationarily reachable cell density is relatively small.

Therefore, the efforts in the development of reactor systems for anaerobic processes for the pro-cessing of organically highly burdened process waste waters are directed to the fact to separate the fIuid dynamics of cells and liquid from one another. That means, by suitable measurei~ the wai~hing out of the cells from the reactor by the process water flow is avoided, or the cells washed out are separated from the outlet and are returned into the reactor.

It is possible by such measures to adjust a higher cell concentration ("bio-catalyst") in the reactor and, thus to increase the velocity of reaction so that also higher concentrated inlet flows may be treated with acceptable residence time (reduction of ~ ,'J``,',~
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f~the residence time of 10 - 20 days to 1 - 2 days~.

For a conversion as complete as possibler however, in the case of a bio-reactor of the agitator vessel S type still relatively long residence periods are necessary. However, if the conversion would be carried out in a tube reactor or in a multistage cascade, consequently with a minimum back mixing of the liquid, then with a shorter residence time and, thus, a s~aller reactor volume an almost complete conversion could be reached. However, in such a reactor the pH value ~ould not be uniform locally; rather same would increase from the inlet value (tha~ means, the-pH value in the process water~ when entering the reactor) along the flow path to the value in the outlet. Thereby:optimum conditions for the micro-organisms would not be guaranteed and, thus, their efficiency of decomposition would be af-fected and hampered.
.

In order to provide, on the one hand, optimum conditions for the conversion regardir.s t}.e F~ va1~
however, on the other hand to reach a high total turn-over, in the case of the process according to the present invention the microbial anaerobic reaction in the conversion of methanogenic substrates (acetogenic and methanogenic phase) is carried out in a two-stage reactor system. In this connection, in the first~stage ~not to be confused with the first phase of the anaerobic decomposition of complex substrates at all, the fer- :
mentative-hydrolytical phase) in a reactor of the agi-tator-vessel type, the main conversion of the methano-genic substances (inter alia acetic acid) is carried out at the optimum pH value of up to approximately 90 %. In a reactor arranged thereafter with the re-sidence time behaviour of a tube reactor or of amultistage cascade, respectively, the remainder of .

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-` 1243471 the substal1ces is converted into biogas up to a de-sired total conversion. By this, the pH value does not change considerably any more. It can be shown that by such a combination a total conversion of, e. y., 99.9 ~ can be reached within a total residence time, which is just 20 % of the residence time, that is necessary for the one-stage conversion in cn agitator-vessel reactor.

5uch a combination is of significance, if total conversions of more than 95 % are required. For instance, that is the case, if the water processed by anaerobic micro-organisms shall not be purified or cannot be purified in a further biological process stage. In other words, this process proposal enables the exten-sive and far-reaching anaerobic purification of the water; however, in contrast to the aerobic purification energy is not consumed (aerationj, but is gaine~ in the form of biogas. In addition, usually a complete conversion as far as possible lS aspired after just for economical reasons.

In order to be in a position to treat anaerobi-cally the process waters occurring e. g. in the DMT
production, a high bacteria concentration must be reached. For example, this~may be secured in that the bacteria are fixed on inert carrier particles, and/or grow on such suspended ca~rier particles. Technical methods for the fixinJ of cells ~immobilization) on particle-shaped carriers are known.

In the two-stage anaerobic process of the present invention as the first stage a mammoth loop reactor (MSR) may be used, as same is shown, for example, in Fig. 11. The carrier particles, covered with anaerobic bacteria (acetogenic and methanogenic cells) are suspenaed in the liquid. By appropriate geometrical -.

', . ~

12~3~t~1 design of the MSR the fluid content thereof -- the suspension -- is circulated and thereby mixed in the stationary operation by the biogas, formed in the fluid.
However, the propulsion o~ the circulation is also possible by condensing and returning of biogas or by the addition of a foreign gas (e. g. nitrogen when starting). However, in the processing of organically highly burdened process water in most cases so much biogas results that such a strong circulation becomes possible, by which the carrier particles, the bio-catalyst, are fluidized without difficulties and the suspension are properly mixed. The configuration of the MSR as a tall, slim reactor having a height-to-diameter ratio larger than 10 supports the self-acting circulation by the biogas.
.

As the second stage advantageously a so-called self-fluidizing whirl cell tower (WZK) is used. Same is characterized in that the through-flow with fluid (suspension), with corresponding selection of the geometry and the number of cells, practically is ef-fected without back mixing. The WZK is a multistage cascade.

Characteristical for the geometrical arrangement of MSR and WZK in the sense of the process of the present invention is that the "active height" of the WZK must be smaller than the height of the circulation system in the MSR, in the case of the MSR with inner circulation, consequently, smaller than the length of the installed guide tube. Then the circulation in the MSR can be utilized, in order to circulate a fluid flow from the MSR from top to bottom through the WZK
and back to the MSR. This fluid flow through the WZK
normally is just a fraction of the fluid flow, cir-culating in the MSR.

.

i243'~'7~

If the MSR is operated in such a manner that the suspension ~lows in the lnner space upwards an~ in the outer space downwards, then the head of the WZK is connected to the outer space of the MSR a lit-tle below the upper guide tube edge by a descending pipeline.
From the bottom of the WZK a likewise descending pipe-line leads to the MSR, if possible into the region of the lower ~eflection at the end of the outer space~
Thereby the same pressure difference, which conveys the fluid through the outer space of the MSR, also acts upon the WZK.

However, it is not the sense of the arrangement to recirculate a suspension partial flow completely;
rather it is intended to return just the particle flow into the MSR. The carrier fluid, which was freed in the MSR by the micro-organisms partly from the orga-nic substrates, shall be further treated in the WZK
to a remaining content as low as possible and then shall be remo~ed from the system. For this serves:the property of the WZK, wherein the fluid, in the parallel flow as well as in the counterflow with the particles, streaming downwards, of the bio-catalyst, is subject to the microbial conversion almost without any back mixing.

In the case of pure parallel flow in the WZK
the suspension flow from the MSR enters with about the same content of particles as in the MSR the head of the WZK. Particles and fluid flow through the WZK
from top to bottom. The particle concentration and the residence time of the particles are adjustable by means of the geometry of the installation le. g.
inclined trays or conical plates). The particles are fluidized by the flow in the cells of the WZK/withu-t supply of external energy. After flowing through an , _ . . . .

- ~.2439L'7~
L/~

appropriate number of cells the substrates are decomposed to a concentration corresponding to the desired total conversion. A part of the 1uid flow now leaves at the outlet stud un~er the hydrostatic pressue of the system. The outlet position may be designed in such a manner that no particles are re-moved into the processed fluid flow. The rest of the fluid flow now streams togèther with the particles as thickened suspension back again into the MSR.

In the case of pure parallel flow consequently a part of the fluid flow, processed in the WZK to the desired conversion, is recirculated with the particles into the MSR. That means, that the fluid flow, which;
streams from the MSR into the WZK, must be greater than the throughput of process water supplied to the MSR and, thus, the volume of the WZK must be selected larger than the throughput and the residence time for the desired conversion would correspond to.
As shown in Fig. 8, that may be bypassed by distribution of the flows and by counterflow of par-ticles and fluid in the lower portion of the WZK.
Thus, for example, it is possible to supply with the suspension, flowing through the line 3, from the MSR
just half of the fluid flow, corresponding to the throughput, to the WZK at the top. A liquid flow streaming through the line 19 and corresponding to the throughput is taken from the MSR without particles at a suitable position (at the height of the mixture level) and is supplied to the WZK at the bottom. The flow is separated with correct adjustment into a counterflow, which now amounts to half the throughput and counterflows the particles in the WZK as well as into a correspondingly large flow streaming through the line 4, which carries out the particles on the bottom of the WZK and conveys them back into the MSR.

.~ ~.

3~7~
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The counterflow is processed in the WZK by the bac-teria, when streaming through the cells and leaves the column together with the likewise processed parallel flow, coming from the top, approximately at the middle of the column, also in this case removal being pos-sible without simultaneous conveving of particles.

By the fact that the biggest portion of the conversion is carri~d out already in the MSR, just little biogas results in the WZK. This biogas flows through the column without influencing the fluid flow upwards and leaves the head thereof into the biogas removal system.

If in the processing of acidic process waters, as, for example, in the case of the DMT process water, the organic acids are present in the form of acid (and not in the form of salts), then it may be necessary even on the basis of proper mixing of the first stage, to neutralize a part of these acids so that in the first stage the optimum pH value results. For that, the process of the present invention may be applied in a modification.

I~, ~or example by soda lye the inlet is neu-tralized so far that in the stationary operation the optimum pH value results in the first stage, then one finds the amount of sodium ions, corresponding to the inlet concentration of sody lye, in the outlet of the second stage, which keeps,in compliance with the protolytic equilibrium for carbon dioxide,hydrogen carbonate ions in the solution. The application of sodium hydroxide for the part neutralization can be reduced considerably, if one leads, according to Fig.
12, the outlet8from the second stage via a ion exchanger st~ge with cation exchange resin in the form of hydrogen.

, ' 12~34'~1 Then the cation exchanger is charged with the sodium ions. In the outlet 34 carbon dioxide and water are formed from the hydrogen ions and the hydrogen car-bonate ions.
If one now regenerates the cation exchanger charged with the sodium ions by the acidic process water of the inlet 35, then the sodium ions are given into the process water, whereby they partially neu-tralize the acids. Therefor the cation exchanger bindshydrogen ions. For the regulation of the pH value in the bio-reactor it is necessary to add just the amount of sodium hydroxide, which cannot be removed by ion exchange from the processed process water flow.
Example 1:

The organic components, utilizable by methane bacteria, of a process water flow of ca.
VL = 10 m3/h from the DMT production (primarily acetic acid, formic acid, methanol), according to the process as described, are converted in a two-stage plant, con-sisting of an MSR and a WZK, to methane and carbon dioxide. Their concentration is, expressed in equi-valents acetic acid, ca. 1 kmol/m3 corresponding toapproximately 60 kgs/m3 acetic acid. The pH value of the process water is approximately 2.4.

In order to reach, in the case of 90 % conversion - 30 in the first stage a pH value of 7, the pH value in the inlet must be adjusted to pH = 4. That is effected in a ion exchange stage b~ regeneration of a cation ex-changer, charged with-sodium ions and subsequent accurate adjustment by soda lye with totally 0.15 kmol Na per m3 process water.

. , . ~

' ' ' .

lZ~34'7~

~ ~ , j , ~ , For a 90 ~ conversion of the acetic acid ln the first stage ~MSR) a residence time of seven hours is necessary. Same results from the reaction kinetics of the utilized bio-catalyst. The bio-catalyst consists of methane bacteria, which immobil~ed on inert carrier particles as an average having a diameter of 0.1 mm, and is suspended in a volume concentration of ca. 24 %
in the MSR.
.

In the MSR approximately 380 m3/h biogas result, which effects the circulation of the suspension around the guide tube. The gas comes out at the top from the first stage. It has a methane content of approximately 50 %.
The MSR has a volume of ca. 82 m3, with a height of ca. 30 m and a diameter of ca. 1.85 m.

For the arrangement of the second stage two variants are possible:

Variant 1:
.

At a height of approximately 27 m, where the sus-pension, to a large extent freed from gas bubbles, flows into the outer space of the MSR, ca. 20 m3/h suspension with a content of ca. 0.1 kmol/m3 acetic acid stream out from the MSR. They are supplied at the head to the second stage, the WZK. The suspension flows through the column from top to bottom, the remainder of the acetic acid being decomposed almost completely apart from a rest content of ca. 0.1 ~ of the original concentration (consequently with 99.9 ~). For this a residence time in the WZK of ca. 1 hour is required.
At the foot of the column, on the one hand, the 1243~'7~

purified process water 10w, free from particles, is taken and removed, and, on the other hand, the thickened suspension flow of ca. 10 m3/h, which contains all the added particles, is separated out. This flow streams back, under the inEluence of the pressure difference, into the lower deflection in the MSR. Then the suspension has approximately double the value of the inlet con-centration (ca. 48 %).

The volume of the WZK is ca. 15.6 m3, with a height of 20 m corresponding to a diameter of ca. 1 m.

If one would carry out the processing of the process water with the conversion of 99.9 ~ with equal kinetic data of the bio-catalyst in one stage (agitator-vessel reactor), then it would be necessary that this reactor has a volume of 436 m3, correspond-ing to a resldence time of 37.5 hours.

Variant 2:
..
As in the case of the variant 1 a suspension flow, however, of a smaller amount, of ca. 3 m3/h of pure liquid (ca. 4 m3/h suspension) streams at the top into the WZK. The liquid is purified in parallel flow with the particles. 10 m3/h pure liquid stream in a pipeline from the first stage in the vicinity of the mixture level, (where the fluid is practically free from particles), from the bottom into the WZK.
3 m3/h of this liquid flow, which also has the acetic acid concentration of ca. 0.1 kmol/m3, corresponding to the 90 % conversion in the MSR, convey the bio-catalyst particles, travelling downwards through the WZK, back into the MSR. The remaining 7 m3/h stream in a counterflow towards the particles in the WZK
upwards, the acetic acid being further decomposed.

" .

~2~3~71 _ ;L _ The removal of 10 m'/h processed, purified liquid, namely of the parallel flow (3 m3/h) and of the counterflow (7 m3/h), is carried out at the position of the WZK, where in the two flows the acetic acld is decomposed to the rest content of 0.001 kmol/m', corresponding to the total conversion of 99.9 ~.

In the case of this variant the residence time in the second stage requires a reactor volume of the WZK of only 10.2-m3 with the same mean concentration of the bio-catalyst as in the case of variant 1.

In each case the processed water flow contains cells, which detached from the carrier and which~may be separated, e. g., in a filtration stage, dissolved reaction products of the bacteria, remainders of the nutrient salts and sodium hydrogen carbonate according to the amount of the sodium ions, added for the pH
ad]ustment (0.15 kmol/m3). The metal cations of the nutrient salts and the sodium ions may be removed from the water flow in the subsequent ion exchange stage.

In this example, apart from the conveying of the process water into the MSR~against 30 m liquid column, no further mechanical energy (drives) is necessary. The purified water may be conveyed practically with the same conveying energy further, due to the fact that the process stage has a very low pressure loss.
Process 2 Fermentation of sacchariferous solutions for the preparation of ethanol.
Similarly to the anaerobic processing of process .

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.
-2a~34t7~

water the two-stage process procedure may be utilized under application of the process of the present in-vention for the fermentation of ~acchariferous solutions for the production of ethanol.
In the fermentation of sugar (e. g. glucose) by yeasts or bacteria mainly ethanol in aqueous solution and carbon dioxide, which disappears in the gaseous phase from the fluid, result as products.Also in the production of ethanol by fermentation of saccharum one is interested in a high conversion, particularly, if large quantities of low concentrated sugar solution shall be fermented. The rest sugar in the outlet of the fermenter means loss in any case.
Such a fermentation is carried out advanta-geously in two stages. If an MSR is used as the ~irst stage, then the carbon dioxide, resulting in the MSR
on the basis of a conversion of,e. g., 90 %, may be utilized for the propulsion of the circulation~ Owing to the circulation now likewise a suspension flow through the second stage, advantageously a WZK, may be circulated. In the second stage the remainder of the dissolved sugar may be further fermented to ethanol almost completely with a reasonable residence time.
: . . .
In order that the micro-organisms in the WZK
or in the second stage, respectively, are separated 30 from the product flow without difficulty, it is of - -advantage to carry out the fermentation of the sugar with immobilized cells (yeasts or bacteria).

However, the process is not limited to the im-mobilization of the cells, just as little as the above mentioned process (the anaerobic processing of process water). It is also possible to work with freely r ,. ' , lZ~3~7~

cells.
suspended/ In this case, the suspension, after flowing through the WZK, may be led through a separating stage, e. g. a sedimentation stage, or a filtration stage, wherein the suspension is concentrated up. The thickened suspension then flows, as described, under the influence of the pressure difference in the MSR back to same, and the product flow streams away under the in f luence of the hydrostatic pressure of the device from the separating stage.

Example:

In the device, consisting of M5R and WZK, a sugar solution (ca. 1.1 kmol/m3, correspondin~ to 200 kgs/m3 glucose) is fermented by yeast immobilized on carrier particles, to ethanol. The sugar solution is supplied to the MSR. With a throughput of ca. 10 m3/h sugar solution, on the basis of a conversion o~
90 ~ in the MSR, 82.8 kgs/m3 ethanol and approximately 400 m3/h carbon dioxide,~which propels the fluid circulation in the MSR, are produced. Due to the higher velocity of reaction of the ethanol fermentation the residence time is shorter Ithan in the example of process f). The volume of the first stage is approx.
25 m3, the volume of the second stage is approx. 3 m3 ~or a total conversion of 99.9 %.

Claims

The embodiments of the invention in which an exclusive property or privilege is claimed are defined as follows:

1. A process for carrying out reactions and mass transfer processes in heterogeneous fluid systems, particularly solid-liquid systems and solid-liquid-gas systems, characterized by the utilization of a reactor with agitator vessel characteristic as first stage and of a reactor having a smaller volume with tube characteristic as second stage, the two reactors being in communication with each other on the basis of a communicating hydraulic system in the region of their upper and lower ends.

2. A process according to claim 1, wherein the reactor with agitator vessel characteristic is a mammoth loop reactor.

3. A process according to claim 1, wherein the reactor with tube characteristic is a whirl cell tower.

4. A process according to one of claims 1, 2 or 3, wherein the reactor with tube characteristic is integrated into the reactor with agitator vessel characteristic.

5. A process according to one of claims 1, 2 or 3, wherein the inset is supplied to the reactor with agitator vessel characteristic through the lower end of the reactor with tube characteristic.

6. A process according to one of claims 1, 2, or 3, wherein the product removal is carried out in a separating device, arranged in the lower connection between the reactors.

7. A process according to claim 1, wherein the reactor with agitator vessel characteristic is in comunication, via a by-pass line, arranged a little below its mixture level, with the lower region of the reactor with tube characteristic.

8. A process according to claim 1, wherein the reactor with agitator vessel characteristic is in communication, via a return flow line, arranged a little above its mixture level with a region of the reactor with tube characteristic, located above the inlet of the upper connecting line into the reactor with tube characteristic.

9. A process according to claim 8, wherein in the return flow line a throttle device is arranged.

10. A process according to claim 9, wherein the throttle device is a throttle valve.

11. A process according to one of claims 1, 2 or 9, wherein two or a plurality of reactors with tube characteristic are associated with a reactor with agitator vessel characteristic.

12. A process according to one of claims 1, 2 or 9, wherein two or a plurality of reactors with agitator vessel characteristic are connected in series with, in each case, at least one associated reactor with tube characteristic.

13. A process according to one of claims 1, 2 or 9, wherein in the lower connecting line between the reactor with agitator vessel characteristic and the reactor with tube characteristic a throttle device or a distributing device, respectively, is arranged.

14. A process according to one of claims 7, 8 or 9, wherein in the by-pass line a bump is arranged, delivering in the direction to the reactor with tube characteristic.

15. A process according to claim 1, wherein said two reactors are in communication by pipeline.