The Canadian Journal of Chemical Engineering Volume 75 Issue 6 1997 [Doi 10.1002%2Fcjce.5450750604] Frans W. J. M. M. Hoeks; Carla Van Wees-Tangerman; Karel Ch. a. -- Stirring as Foam

Stirring as Foam Disruption (SAFD) Technique in Fermentation Processes

FRANS W. J. M. M. HOEKSI,*, CARLA VAN WEES-TMGERMA.@, KURT GASSER3, HENRICUS M. MOMME@, SERGIO SCHMIL? and KAREL ChA.M. LUYBEM

LONZA AG, 3930 Visp, Switzerland

*Kluyver Laboratory for Biotechnology, D e w Universig of Technology, Julianalaan 6 7,2628 BC Dew, The Netherlands

31ngenieurschule Wallis, Route du Rawyl4 7, 1950 Sion, Switzerland

4LONZA Biotec sro, Okruini 134,28161 Koufim, Czech Republic

Foam reduction through stimng was studied in 20 L bioreactors with artificial media and with an actual biotransformation process. For a given stirrer configuration and within a certain range of the broth mass, the foam height was corre- lated with the broth mass, i.e. the distance between upper stirrer and dispersion surface, and with the superficial gas velocity. Increasing the stirrer speed often resulted in reducing the foam height. A mechanistic model was developed for the Stirring As Foam Disruption (SAFD) technique, relating the foam height to the horizontal liquid velocity near the dispersion surface. The model illustrates the general applicability of the SAFD technique and points to foam entrainment as the major mechanism for the foam disruption.

On a etudie la reduction du moussage par Iagitation dans des bioreacteurs de 20 1 avec des milieux artificiels et avec un procede de biotransformation reel. Pour une configuration dagitateur donnee et dans une certaine gamme de masse de bouillon de culture, la hauteur de mousse a ete correlee a la masse de bouillon, soit la distance entre Iagitateur superieur et la surface de dispersion, ainsi qua la vitesse de gaz superficielle. Augmenter la vitesse de lagitateur conduit souvent a la reduction de la hauteur de mousse. Un modele mecanistique a ete mis au point pour la technique de desintkgra- tion de mousse par agitation (SAFD); ce modele relie la hauteur de mousse a la vitesse liquide horizontale pres de la surface de dispersion. Le modele illustre Iapplicabilite generale de la technique SAFD et suggere que Ientrainement de la mousse est le mecanisme essentiel de la desintegration de la mousse.

Keywords: foam, fermentation, mechanical foam control, gadfoam entrainment, multiple impellers.

n many technical processes, foam is an undesired phe- I nomenon. There exist quite a number of mechanical, chemical or thermal ways to destroy foam (Pahl and Franke, 1995). In fermentation processes, foaming is caused mainly by proteins. Concentrations of 1 mg/L suffice to influence foaming (Prins and vant Riet, 1987). Foam takes up space and therefore reduces the effective production volume. Secondly, the gas-outlet filters of bioreactors can be blocked or get wet due to foam overflow, risking loss of a fermentation run. This risk could be minimized by operation with a large head space. However, this reduces bioreactor output. Accumulation of substrates and biomass in the foam also reduces productivity (Schugerl, 1985). The enrichment of cells in the foam might cause autolysis of the cells releasing proteins and enhancing the interaction of surfactants with dissolved proteins resulting in more foam. Consequently, reduction of the foam layer in bioprocesses is a must. There are several ways to do this: 1) Addition of an anti-foam agent. The addition of (large amounts of) anti-foam is less desirable, because it reduces gas hold-up and therefore oxygen transfer (Lee et al., 1993, Yasukawa et al., 1991a) and may have negative effects on the purification process after the biotransformation.

*Author to whom correspondence should be addressed. E-mail address: [email protected]

2) A rotating disc foam breaker mounted in the head space of the bioreactor overcomes these disadvantages (Ohkawa et al., 1984, 1987; Yasukawa et al., 1991b). However, for existing large scale bioreactors installing a mechanical foam breaker can be difficult or too expensive (Lee et al., 1993). 3) Installing a conical shape draft tube into a bioreactor has been proposed (Schubert et al., 1993). Inside the draf? tube, the liquid is drawn down by a hydrofoil Lightnin A3 15 impeller and pumped into the annulus. From the annulus the liquid flows over the top of the draft tube into this tube. The proposed mechanism for foam suppression is essentially caused by the overflowing liquid curtain. It would be interesting to have experiences on scale-up of this foam disruption mechanism revealed. Furthermore, comparisons of effective production volumes should be made, because the overflow implies that the dispersion level in the draf? tube is lower than in the annulus.

This paper proposes to reduce the foam layer on the broth through stirring. A statistical analysis of a series of biotransformation experiments at Lornas pilot plant revealed a correlation between anti-foam consumption and stirrer speed. The purpose of the work presented in this paper was to devise experiments, providing evidence that foam can be disrupted by stirring and insight on how this is achieved. The observation that stirring can reduce foaming has been made before (Ohkawa et al., 1984, Yasukawa et al., 1991b), but the authors did not elaborate how this was achieved. On the other hand, it has been stated that the foam layer is not influenced by the stirrer (Schugerl, 1985).

1018 THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING. VOLUME 75, DECEMBER, 1997

TABLE 1 Design, Dimensions and Hydrodynamic Data of Upper Stirrers (Data for the pump capacity coefficient (N ) are taken from Nienow (1990), except for the pitched blade stirrers, for which N, was given by the manufacturer TECHMIX sro YBrno, Czech Republic))

D P, at P,, at Stirrer (swept) W L X No. of 500 rpm 800 rpm type mm mm mm mm blades Po W W

120 24 30 4.5 6 0.72 4.3 62 254 95 19 23.8 2 6 0.72 4.9 22 90

125 60 44 2 4 0.73 0.75 13 54 PBD,,,*** 120 24 44 2 6 0.9 1.51 22 89

*Rushton turbine. **Hydrofoil Lightnin A3 15. ***Pitched blade stirrer pumping downward.

N4

RT120*

PBD,,* ** 95 24 32 2 6 0.9 1.91 8.6 35

0 mm

c1

47 mm .-.

47 mm

Figure I - Schematic presentation of the bioreactor(s) used in the SAFD experiments.

When studying the effects of stirring on foam disruption, not only the stirrer speed, but also the stirrer configuration should be the subjects of investigation. Therefore, a small variety of stirrers was tested. Rushton turbines are still stan- dard in industry, but alternatives such as pitched blade stirrers or newer designs, e.g. hydrofoil impellers, wer6 evaluated as well (Joshi et al., 1982, Nienow et al., 1993). Upper pitched blade stirrers were used in the downward pumping mode because on the pilot scale we observed that upward pumping pitched blade stirrers could push the foam up against the tank wall into the gas outlet of the fermenter.

When devising the experiments on SAFD, the media question proved to be crucial. Testing foam disruption through stirring in an actual bioprocess has the advantage of a high relevance of the work. However, running a bioprocess is quite labour intensive with respect to the aim of the foam disruption studies. The labour intensity can be reduced by using artificial foaming media allowing results to be obtained more efficiently. Artificial foaming media are described in the literature and can be used for comparison, but might introduce artefacts. Therefore, both an actual bioprocess and artificial media were used to study SAFD.

The ultimate goal of the work presented in this paper is, to propose a bioreactor design enabling an enhanced working volume as a result of adequate foam disruption through stirring. Because stirring influences the gas hold-up of the broth and, therefore, the working volume, changes in gas hold-up should be evaluated when applying the SAFD technique.

Materials and methods

The experiments were carried out in bioreactors with a total volume of 20 L with inner diameters of 195 and 200 mm (MBR, Wetzikon, Switzerland). The bioreactors had 4 baffles with a width of 19 mm. A ring sparger below the bottom stirrer was used for air supply. The bioreactors were equipped with two or three stirrers. See Figure 1 for the dimensions of the bioreactors. In Table 1 the geometry and power draw of the stirrers, used in the hghest position are given: i.e. a 6-blade 45 pitched blade stirrer pumping downward (PBD), Rushton turbines (RT) both stirrer types with 95 and 120 mm diameter (subscripts 95 and ,20), and a hydrofoil Lightnin A3 15 (HF). A hollow blade stirrer, type Chemineer with 6 blades (HB), was mostly used as a bottom stirrer, because it has been demonstrated that the bottom impeller determines the gas hold-up (Chiampo et al., 1991) and therefore in principle also the foaming character. Table 2 gives the stirrer configurations tested. These configurations were chosen for practical reasons, such as commer- cial availability of the stirrers.

All series of experiments were carried out in duplicate, i.e. each stirrer configuration was tested in 2 biotransformations. The model media were prepared twice for each series of experiments with one stirrer configuration.

The fed-batch L -carnithe biotramformations were carried out with mineral medium as described before (Hoeks, 1991; Hoeks et al., 1996). After the biomass growth phase, the biotransformation of y-butyrobetaine into L-carnitine was carried out. During this biotransformation phase, foam disruption through stirring was studied.

The compositions of the model media are given in Table 3. The Combinations of stirrer configuration and media composition are given in Table 2. When the stirrer configuration was altered, the medium was made afresh, because surfactants from the skin influenced the experimental results. Moreover, rubber gloves were used for the manipu- lations. In between experiments, the bioreactors were cleaned with deionized water and ethanol.

Lee et al. (1993) related the foam height to the superficial gas velocity. As a rule, the superficial gas velocity increases when scaling up. In order to obtain data relevant for large scale, the experiments on the laboratory scale were carried out at superficial gas velocities which can also be found on large scale, i.e. 0.0065 and 0.013 m/s. This corresponds with a gas flow of 0.00020 or 0.00040 m3/s, respectively, or 10 or 20 L(STP)/min, f 2.5%, respectively, depending on the diameter of the 20 L bioreactor used.

THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997 1019

TABLE 2 Stirrer Configurations Tested (The 6-blade 45" pitched blade stirrer pumping upward (PBLI) had the same dimensions as the downward pumping impeller (PBD). The hollow blade agitator (HB) of the Chemineer type CD-6 had 6 blades. The media are given in Table 3)

Upper Middle Bottom Configuration stirrer stirrer stirrer Media tested

- HB I, I1 HB L-c, I HB or RT,,,

- HB I, I1

RT95 RT95 RT95 RT95 - RT95 RT95

R7-I20 RTl20

HB I PBD95 PBD95

HB I PBD95 PBD95 - PBD95 PBD95

HB L-c, I PBDl20 PBDl20

HB L-c, I PBDl2O - PBDl20 PBDl20 PBD120

HB I, I1 PBUl20 PBDl20 PBDI20 - PB"l20

HF-HF HF HF HB L-c, I

L-c, I, 11,111, 1v -

-

- HF HF

HB - HB I -

TABLE 3 Composition of the Model Media Used (Concentrations in deionised water in % (w/w). The description of the biotransfonnation medium

is given by Hoeks, 1991 and Hoeks et al., 1996) Medium Tween 40 Triton X-100 NaCl

I 0.0002 11 0.0002 4 111 0.00063 IV 0.00063 4 L-c biotransformation

The influence of the stirrer speed on the equilibrium height of the foam layer during the biotransfomation was studied by increasing the stirrer speed by steps of 100 from 500 to 800 rpm. The dispersion volume and the height of the foam layer were measured at each speed. To study the effect of the proximity of the upper stirrer to the dispersion level, 200 f 10 g of broth were taken out of the bioreactor and the above increase in stirrer speed was repeated. The dispersion volume was read from a litre scale fmed on the wall of the bioreactor midway between two baffles with an accuracy of f 50 mL. The height of the foam layer was measured with a scale with an accuracy o f f 0.5 cm. After each time taking 200 g of broth out, the substrate feed rates were reduced to keep the specific substrate feed rates at the same level as before taking the broth out. Measurements of the L-carnitine concentration showed that the biotransformation performance was normal (data not presented).

The exact L-camitine broth mass could only be determined after each biotransformation due to the requirement for aseptic operation and due to technical limitations. At the end of each biotransformation, approximately 750 g less was present than calculated from reduction of broth mass and sample taking. This loss could be explained by evaporation during the biotransformation. The rate of evaporation derived from the above loss in broth mass was used to correct the calculated broth weight during the experiments.

The experiments with the model media were carried out at 500 and 800 rpm only. Considerably less work was involved in the preparation of the experiments with the model media and, therefore, more stirrer configurations could be studied.

Because the experiments with the model media were short in comparison with the biotransformation, the loss on evaporation was negligible. Liquid masses were determined directly by weighing without any corrections.

The gas hold-up cG was calculated from the difference between dispersion volume and broth mass, assuming a liquid density of 1000 kg/m3 for the artificial media without NaCl or using the actual density of 1030 kg/m3 for the broth and the artificial media with NaCl. The filtrate of the L-carnitine broth contained 100 f 10 mgkg protein. The L-carnitine broth had a surface tension of 0.04 N/m.

Modelling

DEVELOPMENT OF THE MODEL FOR THE SAFD TECHNIQUE

The mechanism for foam disruption through stirring could be that the foam at the dispersion level is being drawn into the liquid similar to the phenomena of gas entrainment from the head space into the liquid. This "foam entrainment" has been suggested after observations using a mechanical foam breaker and varying the stirrer speed (Yasukawa et al., 1991b). Lee et al. (1993) assumed that the foam is formed at the dispersion level of aerated vessels at a speed equal to the superficial gas velocity vs or even hgher. If the foam is rising at this speed and if the%oam disruption through stirring would only be caused by foam entrainment, the foam entrainment must happen at a superficial velocity vsg (downward), if the foam height is to be zero. However, thls would lead to the false conclusion that the net airflow at the dispersion level is zero. Thus, foam entrainment alone cannot be the only reason for foam disruption through stirring.

The mechanism of mechanical foam disruption is considered to be breaking up of the liquid lamellae between the gas compartments in the foam by a certain stress on these lamellae (Pahl and Franke, 1995). In this paper, it is proposed that the liquid lamellae between the gas compartments in the foam at the dispersion level can be disrupted by a stress caused by the liquid flow at the dispersion level as a conse- quence of stirring.

Since both foam entrainment and the stress on the liquid lamellae causing foam disruption have the liquid flow at the dispersion level as the common denominator, the basis of the model for stirring as foam disruption (SAFD) technique was to define a parameter representing the liquid flow at the dispersion level. If this representation is adequate and if the above hypothesis is correct, a correlation between the liquid flow and the equilibrium foam height should be found.

The commonly accepted flow pattern generated by a radial pumping stirrer consists of an upward flowing component and a downward flowing component starting fiom the impeller

1020 THE CANADlAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997

plane at the tank wall (Josh et al., 1982, see Figure 2). In this paper it is assumed that only the upward component of the flow generated by the upper stirrer is relevant for SAFD. For axial downward pumping stirrers, the liquid flow was considered to consist of only one loop around the impeller (see Figure 2). Axial pumping sthers form a rapid recircu- lation loop around the tips of the impeller blades when aerated (Manikowski et al., 1994).

MATHEMATICAL ELABORATION OF THE SAFD TECHNIQUE

A cylinder with a diameter of half the tank diameter (0.5T) from the middle of the upper stirrer to the gas-liquid dispersion surface is defined (see Figure 2). For the mathematical elaboration of the model for axial pumping stirrers, it is assumed that the flow direction of the liquid is horizontal from the wall to the axis over the whole height of this cylinder. Furthermore, it is assumed that the liquid velocity is constant over the height of the cylinder. Consequently, the velocity vL dl of the above defined liquid flow at the position of the cylhder wall is calculated by dividing the stirrer discharge flow under gassed conditions, QL,g, by the vertical cylinder surface A, (see Figure 2). In formula:

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . (1) QLX vL.dl = 1

The scale independent parameter vL,dl is chosen to represent the horizontal liquid velocity near the dispersion level and shall be related to the equilibrium foam height.

For radial pumping stirrers it is assumed that half the total liquid flow generated by the Rushton turbine flows in an upward loop and the other half in a downward loop (see Figure 2). Therefore, the stirrer discharge flow QL,g for the calculation of the liquid velocity vLdl in the upper loop is divided by two ( Q L , 4 2 ) . Note that for radial pumping stirrers, the flow from wall to axis above the stirrer can only exist in the upper part of the upward loop. In the lower part of the upward loop the flow is from the axis to the wall (see Figure 2). Again, a cylinder with a diameter of half the tank diameter (0.5T) from the middle of the upper stirrer to the gas-liquid dispersion surface is defined. For the mathematical elaboration of the model for radial pumping stirrers, it is assumed that the flow direction of the liquid is horizontal from the wall to the axis only in the upper half of this cylinder. Furthermore, it is assumed that the liquid velocity is constant over the height of the upper half of the cylinder. Thus, for the calculation of the liquid velocity VL,d[, the vertical cylinder surface A, must be divided by two for radial pumping sthers (AJ2). Consequently, Equation (1) is valid for radial pumping stirrers as well, because both QL,g and A, are divided by two.

The height of the above defined cylinder is calculated by taking the difference between the filling volume at the dispersion level ( v d ) and the filling volume at the level of the middle of the upper stirrer (V, ) and dividing it by (x/4)P. The vertical surface area of the cylinder is xT/2 times this height (see Figure 2):

. . . . . . . . . . . . . . . . . . . . . . . . (2) (vd - 5) Ac = zOST (n / 4)T2 The pumping capacity of a stirrer depends on the dimen-

sionless pump capacity coefficient (Nq), the stirrer speed

I

T Figure 2 - Schematic presentation of upper stirrer with model cylinder and flow patterns of axial and radial pumping stirrers. V, indicates the filling volume in the bioreactor at the level of the middle of the upper impeller. Vd indicates the filling volume in the bioreactor at the dispersion level.

(N) and the diameter of the stirrer (0) and is given by the following Equation (Oldshue, 1983):

QL,, = Nq N D 3 . . . . . . . . . . . . . . . . . . . . . . . . . . . . . (3) Equation (3) does not take the entrained flow into account

(Oldshue, 1983). The flow generated by impellers is reduced under gassed

conditions by the ratio of gassed to ungassed power draw of the stirrer to the power 0.34 (Rousar and Van den Akker, 1994).

(4) . . . . . . . . . . . . . . . . . . . . . . . . . Q ~ ~ = ( p y ) pg 0.34 QL,u

The ratio of gassed to ungassed power draw is taken from graphs of Pg/Pu against gas flow number FIG found in literature (see references below). The gas flow number, is calculated as follows:

Usually, the correlations between FIG and P /P, are given for single impeller systems. Hudcova et al. ( h 8 9 ) and Smith et al. (1987) have shown that these correlations are also valid for upper impellers in multi-impeller systems at low values of FIG and high stirrer speeds, which is the case in the studies presented here.

Wannoeskerken et al. (1984) gave a Figure in which P IP,, against the gas flow number is plotted for a pitched bfade stirrer with 6 blades pumping downward. The P /P, values determined at gas flow numbers of 0.01 and O.Ofare used for a linear interpolation.

P . . . . . . . . . . . . . . . . . . . . . . . 2 = - 1 7 . 5 F l ~ + 1 . 1 2 5 (6)

p,


The plot of PgIP, against the gas flow number for the hydrofoil Lightnm A315 is taken from Balmer et al. (1987) for the hydrofoil Prochem Maxflo and also linearised. Nienow (1 990) stated that the hydrodynamics of both hydrofoil types are very similar.

(7) P 2=-11Flc+0.73 . . . . . . . . . . . . . . . . . . . . . . . . . . pu

PgIPu for the Rushton turbine can be calculated with the following Equation given by Joshi et al. (1982):

B = o . 1 P ( - gr25[ N2D4 )-I5 . . . . . . . . . . . . p, gwV2l3

The ungassed power draw at 500 and 800 rpm (P,) and the power number (Po) of the various stirrers are listed in Table 1. The ungassed power draw follows from (Joshi et al., 1982):

. . . . . . . . . . . . . . . . . . . . . . . . . . . P, = p Po N3 D5 (9)

Power number data (Po) are calculated from (Bujalski et al.,

Rushton turbine: 1986a).

. . . . . . . . . . . . . . . . . . . . . Po = 2.5 ( X / D ) ~ . io.065 (10) in which T is the tank diameter in m.

PBD, 6 blades, WID = 0.2:

Po = 0.78 (x/D)-O.I4 (D/T)-O. . . . . . . . . . . . . . . . . . (1 1)

Bujalski (1986b) argued that the power number increases proportionally to the ratio of the blade width and the impeller diameter (WID). The power number of the 95 mm pitched blade stirrer is adjusted accordingly. Bujalski (1986b) found that Equation (1 1) is valid for single downward pumping stirrers at DIT < 0.5 and for upward pumping stirrers with higher D/Tratios. In this paper, with multiple impeller systems and DIT not deviating too much from 0.5, it is assumed that Equation (1 1) can still be used. The hydrodynamic data for the hydrofoil Lightnin A3 15 are taken from Nienow (1990).

COMPARISON WITH LITERATURE

A literature study did not reveal any publication on the use of stirrers with the purpose of foam reduction. However, when studying a rotating disk mechanical foam breaker, effects of stirrer speed on foaming have been observed (Ohkawa et al., 1984). Ohkawa et al. (1984) also observed effects of liquid volumes on foaming, which are covered in the above model by reducing the distance between the upper stirrer and the dispersion level, thus increasing the liquid velocity at the dispersion surface. In a publication of Bakker and Frijlink (1989), drawing down and dispersing floating solids (polystyrene spheres) were presented. They concluded that upward pumping impellers close to the surface are most efficient for drawing floating solids into the gas-liquid dispersion. It was found that the creation of a vortex, which is advantageous for drawing down the floating solids, is not easy under aerated conditions (Bakker and Frijlink, 1989).

Veljkovic et al. (1991) reported on the surface aeration of sparged and agitated vessels, which has long been under- stood as gas entrainment from the head space into the liquid. Veljkovic et al. (1991) found a correlation for the minimum stirrer speed N, required for the onset of gas entrainment in unaerated vessels with one Rushton turbine (DIT = 0.33):

. . . . . . . . . . . . . . . . . . . . . . . . . . N,D = 0.732 m/s (12)

Veljkovic et al. (1991) found the following correlation for Rushton turbines of Dierendonck et al. (1971) fitting their experimental data for unsparged conditions adequately:

. . . . . . . . . . . . . . N, = 1.55( T/D2)( h/T)( crgIpL ) 0.25 (13) Equation (13) can be used for every geometry. Note

that N, is linear dependent on the distance h between upper stirrer and liquid surface, which is consistent with the model presented here. For the geometries presented here, the onset of gas entrainment for the 95 and 120 mm Rushton turbines under ungassed conditions would be around 100-200 rpm according to Equation (13). This is far below the stirrer speeds used in the experiments reported here. In other words, in the experiments reported here the gas entrainment from above the broth surface must have been high.

Interestingly, Equations (12) and (1 3) imply that NsD is a constant for a given geometry and a given liquid. As can be derived from Equations ( 1 x 3 ) for a given geometry:

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . v ~ , ~ , a ND (14) Consequently, the approach of relating the foam height to

the liquid velocity at the dispersion surface is consistent with the considerations concerning gas entrainment from the surface. Furthermore, Equation (14) suggests that the tip speed is the parameter to be kept constant when scaling up the SAFD technique.

Results and discussion

FOAM HEIGHT

For all experiments the equilibrium height of the foam layer was plotted against the L-carnitine broth mass with the superficial gas velocity as a second parameter. For each stirrer configuration and each stirrer speed, separate plots were made. Striking similarities were observed. Therefore, a selec- tion of the data for graphical presentation was made.

In Figure 3 the foam height is depicted as a function of the broth mass for the 500 and 800 rpm experiments using the 120 mm Rushton turbine as upper stirrer in the L-camitine biotransformation. Figure 3 shows that the foam height is reduced by stirring faster within a certain range of broth mass. Furthermore, increasing the superficial gas velocity gave an increase in the height of the foam layer. Experiments with 600 and 700 rpm resulted in foam heights between the values shown in Figure 3 (data not shown). The RT,,-RT9, combination, the PBD120-PBD,20 combination and the 2 hydrofoil stirrers show comparable relationships between foam height and broth mass for a given airflow and given stirrer speed. Only the data of the dual Rushton RT95-RT95 are presented (see Figure 4). Again the foam height can be reduced by stirring faster. At 800 rpm, the RTgrRTg5 combination seemed slightly more effective in foam reduction than the PBD,,o-PBD,20 and the dual hydrofoil combination, particularly at higher broth mass (data not shown).

1022 THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 7 5 , DECEMBER, 1997

800 rprn

30

25

1 20 Y

15

L I 10

-

--

._

--

--

800 rpm

c I l o "I

8 9 10 1 1 12 13 14 broth mass [kgl

Figure 3 - Foam height as a function of broth mass and superficial gas velocity for the 120 mm Rushton turbine (RT,20) as upper stirrer at 500 and 800 rpm. Duplicate L-camitine biotransfoma- tions. The open and the closed symbols distinguish between the data of either the one or the other biotransformation. 0 +: vsg = 0.0065 d s , 500 rpm A A: vsg = 0.013 d s , 500 rpm 0 W: vsg = 0.0065 d s , 800 rpm 0 0: vsg = 0.013 m/s, 800 rpm

Comparison of all the upper stirrers at the highest superficial gas velocity and a given stirrer speed shows that the large Rushton turbine of 120 mm diameter resulted in the lowest foam height for a given broth mass (compare Figures 3 and 4). During the experiments reported here, it was observed that the 120 mm Rushton turbine created the largest vortex and had the best foam reducing properties. This corresponds well with the observations by Frijlink and Bakker (1989) on drawing down of floating solids. Tanaka and I m i (1987) found that larger impellers are more effective concerning gas entrainment.

Of course, the power draw of the 120 mm Rushton is much larger than those of the other stirrers for a given stirrer speed. However, the Rushton turbine of 120 mm diameter draws less power at 500 rpm than the 95 mm Rushton turbine at 800 rpm (see Table 1). Still, the foam height at a given broth mass and a given superficial gas velocity was considerably lower for the 120 mm Rushton turbine at 500 rpm than with the 95 mm Rushton turbine at 800 rpm (compare Figures 3 and 4). Consequently, for a given power draw large stirrers at a low stirrer speed have to be preferred (see below). The hydrofoil has a lower ungassed power draw than the RTss and the PBD,,, (see Table 1). A hydrofoil of 140 mm would have an ungassed power draw of 23 W at 500 rpm, comparable to the power draw of the RTs5 and PBD,,,. Therefore, the foam reducing properties of a 140 mm hydrofoil should be better than those of a 95 mm Rushton turbine or a 120 mm PBD at equal stirrer speed.

0 0 .

0 o o o o 8

0

500 rpm .* o A o

A A 0 0 0

5 t 0 4 I 8 9 10 11 12 13 14

broth mass Ikgl

Figure 4 - Foam height as a function of broth mass and superficial gas velocity using two 95 mm Rushton turbines (RT9> - RT?,) as upper stirrers at 500 and 800 rpm. Duplicate L-camitine biotransformations, For symbols see Figure 3.

Consequently, a hydrofoil impeller could be suggested for retrofitting a bioreactor in order to apply or improve SAFD.

When using the model media, the phenomena concerning foam disruption as described above for the biotransfonnation system were to a large extent observed as well. Not surpris- ingly, the experiments with the model media were much more reproducible in comparison with the biotransfonnation experiments (see figure 9 for example). Figure 3 at 500 rpm illustrates the variation introduced by the biological system: The closed triangles in figure 3 seem to form 2 lines, which are caused by carrying out the SAFD experiments at 2 different time intervals within one biotransformation.

FOAM MAP FOR SAFD

The experiments on stirring as foam disruption technique show 3 "foam regimes": 1) In practically all systems studied, there seemed to be a certain maximum broth mass below which there was hardly any foam present. In other words: It appears to be possible to create enough stress on the liquid lamellae and enough foam entrainment by stirring to disrupt the foam of a nor- mally foaming system completely. This so called maximum foam free broth mass depends on the superficial gas velocity, the stirrer configuration, the stirrer speed and the medium (see Table 4). Generally speaking, increasing the stirrer speed from 500 to 800 rpm gave an increase in the maximum foam free broth mass of 4 to 20%. However, there were

THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997 I023

exceptions where the opposite occurred, e.g. with the PBD120-PBU120 combination, which was the worst foam reducing stirrer combination of all. 2) At a broth mass higher than the maximum foam free broth mass, there was in many cases a range of broth mass in which the foam height was almost linear dependent on the broth mass for a given superficial gas velocity and a given stirrer speed (see Figures 3 and 4). Doubling the airflow in this range of broth mass resulted in a foam layer, which was approximately twice as high (see Figures 3 and 4). At constant broth mass in the range of foam regime 2 and at constant airflow, increasing the stirrer speed from 500 to 800 rpm gave a reduction in foam height of 20 to 50%. 3) At a higher broth mass still, in some cases there seemed to be a height of the foam layer which was independent of the broth mass, or the foaming could not be kept under control. The observation that there seemed to be a constant height of the foam at higher broth mass was more profound with the model media. When using Medium I and only stirring with the hollow blade impeller at the bottom of the tank, the foam height was 7-8 cm, practically independent of the superficial gas velocity. This observation is not consistent with Bikerman's theory on foaming, but has been observed before (Lee et al., 1993). Consequently, elaborating SAFD only makes sense if the equilibrium foam height is in the range, in which the foam height is influenced by stirring, i.e. regimes 1 and 2. This equilibrium is system dependent. Note: The equilibrium foam height was 2 to 3 times higher for the biotransformation system as compared to the model media.

The observations described under foam regime 2 are consistent with Bikerman's theory concerning the dynamic equilibrium of foam, which Lee et al. (1993) redefined as a constant ratio between foam height and superficial gas velocity. Through stirring this dynamic equilibrium can be influenced, but at too high distances between upper stirrer and dispersion level, foam reduction through stirring does not occur anymore (foam regime 3). This compares well with observations on gas entrainment: Tanaka and Izumi (1987) showed that reducing the liquid height in a stirred tank facilitated gas entrainment strongly. Thus, at too high distances between upper stirrer and dispersion level, not stirring but other phenomena determine the foam height, such as redistribution of surfactants, secondary foam formation, foam mass, etc. (Pahl and Franke, 1995).

Interesting is the comparison between configurations with one and with two upper stirrers of the same type and size. There was a tendency that one stirrer was more effective than two. This is best shown by the comparison of the maximum foam free broth mass, which was up to 15% higher with one stirrer (see Table 4a). If the flow generated by the upper stirrer is the cause of foam disruption, a reduction in foam disruption ability caused by the middle stirrer in close proximity points at an interference of the flows generated by the upper and the middle impeller. As Chiampo et al. (1993) have demonstrated, this interference is strong at low impeller spacings, as is the case here with 3 impellers, and increases with increasing superficial gas velocity. The interference of the flows may also explain why the PBD,20-PBUl,o combination did not have good foam disruption properties.

Analogous to the comparison of the Rushtons of different diameters, 95 mm pitched blade stirrers were less effective in foam disruption than 120 mm pitched blade stirrers (see Table 4). In fact, stirring with 95 mm pitched blade stirrers influenced the foam height only marginally if at all.

F Ei C 3

1024 THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, I997

TABLE 4B Maximum Foam Free Broth Weight (foam height < 0.5 cm) for

the 120 mm Rushton Turbine as a Function of Stirrer Speed, Superficial Gas Velocity and Medium (The maximum foam free

broth mass was determined by extrapolation of the linear relationship of foam against broth mass)

Maximum foam free broth mass (kg) 0.0065 m/s 0.0065 m/s 0.013 m/s 0.013 m/s

Medium 500rpm 800rpm 500rpm 800rpm I 12.5 11.4 11.5 10.8 I* 12.3 11.7 10.7 10.8 I1 9.5 10 8.4 9.0 111 10 > I I ? 9 9.9 IV 8.8 9.9 8.0 8.7 L-c* 11.5 13.5 10.8 12.6

+RT,20 as bottom stirrer, otherwise HB as bottom stirrer.

CORRELATION BETWEEN THE CALCULATED LIQUID VELOCITY AND THE HEIGHT OF THE FOAM LAYER

For each measurement of the foam height, the liquid velocity near the dispersion level was calculated according to the model described in this paper (Equation (1) - (8)). The foam height was then depicted as a function of the calculated liquid velocity vL,dl for each set of stirrer configuration and medium, irrespective of the superficial gas velocity, of the stirrer speed or of the broth mass. In Figure 5 , the foam height is given as a function of the calculated liquid velocity near the dispersion level for the 95 mm dual Rushton combination and for the 120 mm Rushton for the L -camitine biotransformations. Although the foam heights measured were rather different for these stirrers (see Figures 3 and 4), plotting them as a function of vL,dl results in strik- ingly similar pictures. Thus, the mechanistic model relating the liquid velocity near the dispersion level to the foam height appears adequate to describe foam disruption through stirring. Above a calculated velocity of approximately 0.28 m / s practically no foam is present anymore (see Figure 5). Note the outliers of one biotransfomtion (see also Figure 4).

A correlation in the form of a hyperbolic relationship between the liquid velocity near the dispersion level (v& and the foam height seems to be an adequate mathematical description:

. . . . . . . . . . . . . . . . . . . . . . . . . . . . (15) h f = + - P vL,dl

Assume that the liquid velocity has a certain value vL,dr,o above which the foam is practically absent, e.g. a value of 0.28 m / s as mentioned above for the Rushton turbine. In principle this corresponds with the maximum foam free broth mass (see Table 4) and can be considered as a bound- ary condition for Equation (15):

. . . . . . . . . . . . . . . . . . . . . . hr = 0 for v ~ , ~ , = vL,dl,o (16)

Thus, Equation (1 5 ) can be rewritten as:

(1 7) vL,dl,O

vL.dl hf =a(l-- ) . . . . . . . . . . . . . . . . . . . . . . . . .

30 T

H 3 hf = -1 4.98 * (1 - 0.26 / vL,~J 15 e

120 mm Rushton

I

%+

+ +

++

h, = -16.59 * (1 - 0.28 / v L ~ J p++ + +

++ ; f + + + +

k+++++ + 05 mm Rushton h + +

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

vwI [m/sl

Figure 5 - Foam height as a function of the calculated liquid velocity for two 95 mm Rushton turbines (RT,, - RT94 as upper stirrers and for one 120 mm Rushton turbine (RT,,,) as upper stirrer. Data from all experiments, i.e. 500,600,700,800 rpm and 0.0065 and 0.01 3 m/s. Duplicate biotransformations. The Equation and the line represent the correlation of the two parameters according to Equation (1 7).

VL,dl,o was estimated graphically for each stirrer combination. a was determined by linear regression of the equation:

y = m, with hf = y and

. . . . . . . . . . . . . . . . . . . . . . . . . . . (18) vL,dl,O ) vL,dl

z=(l--

Outliers were rejected by graphical judgement and a was determined according to the least square method with Windows Excel 5.0. Foam heights at too low liquid velocities, i.e. too high distances between dispersion level and upper stirrer, were not used for regression, because foam was not being disrupted anymore by stirring (foam regime 3). The correlations according to Equation (1 8) for the sets of stirrer speed and superficial gas velocity for each stirrer turned out to be parallel lines (see the example in Figure 6). Still one best fitting correlation for each stirrer configuration was for- mulated. The parameter values of the correlations are given in Table 5. Note, that the R-squared values of the above linear regression are low. The main reason for this is the formula- tion of one correlation for all experimental conditions for a given stirrer configuration and medium despite the above observation on the parallel lines. This shows that the presented mechanistic model is not 100% adequate to describe the dynamic equilibrium of the foam as a function of stirring only. This conclusion is consistent with the observation of Lee et al. (1993), that the equilibrium is system dependent.

The correlations according to Equation (17) for the L - camitine biotransformations are graphically represented in Figure 7. Comparison of these correlations shows in the first


d

+ -1 .o -0.5 0.0 0.5

Figure 6 - Graphical presentation of Equation (18) for the 120 mm Rushton turbine as upper stirrer and the hollow blade as bottom stirrer in medium I with 4 parameter sets. For symbols see Figure 3.

2

TABLE 5 Correlation Coefficients for Equation (1 7) for the Stirrer Configurations and Media Tested, and the Corresponding

R-Squared Values Configuration Medium 01 [cm] v ~ , ~ , , ~ [ds] R-squared

I I1 I L-c I I*** I1 Ill IV L-C*** I I I I L-C I L-C I II** I L-C

-7.80 -1 1.92 -5.23

-16.59 -12.40 -10.75 -13.32 -97.69 -27.83 -14.98 (-0.4 I )* (-1.63)* -5.55 -3.43

-41.24 -3.21 -5.94 -3.84 -6.09 4 . 4 5

-29.48

0.16 0.72 0.15 0.52 0.18 0.73 0.28 0.38 0.20 0.28 0.2 1 0.53 0.20 0.12 0.20 0.84 0.2 1 0.23 0.26 0.20

(2.0)* (0.70)* (1 .O)* (0.06)* 0.72 0.52

-0.75 0.35 0.85 0.55 0.75 0.59 2.20 0.40 0.55 0.38 0.44 0.37 0.60 0.04 0.58 0.76

~~

*no v ~ , ~ , , ~ for these stirrer configurations. **800 rpm only. ***RT,20 as bottom stirrer.

place that the relationship between the calculated velocity v ~ , ~ , and the foam height depends on the type of the stirrer (see Figure 7). Assuming that the medium properties in all biotransformations were more or less the same, this means that the simplified flow model presented in this paper does not describe the differences between stirrers in flow characteristics, which are relevant for SAFD, adequately. For example, the entrained flow is stirrer type dependent (Joshi et al., 1982) and explains at least partly why the parameters of the correlations according to Equation (17) are stirrer type dependent. The entrained flow of a Rushton turbine is largest and almost equal to the pumping capacity calculated with Equation (3) (Joshi et al., 1982). However, taking the entrained flow into account would still give a dependency of

I 10 e

-PBD,a - PBDla

- PBD,,, - PBU,, HF - HF I.-

0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0

VLdl [ d s l

Figure 7 - Graphical presentation of the correlations of Table 5 for the biotransfonnation experiments with the 5 stirrers configurations tested (see table 2).

the correlations according to Equation (17) on the stirrer type (data not shown). The uneven distribution of the flow is also stirrer type dependent (Joshi et al., 1982) and con- tributes to the stirrer type dependency of the parameters of Equation (1 7).

The correlations between foam height and liquid velocity are practically the same for the Rushton turbines of 95 and 120 mm (see Figures 5 and 7). This means that the chosen model, which describes the phenomenon of foam disruption through stirring, is geometry independent! Because the correlations for the Rushtons are practically the same, the model provides also an adequate explanation for the better foam reducing properties of the 120 mm Rushton as compared to the 95 mm Rushton at equal power draw. For constant power draw, it can be derived from Equation (9):

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . N oc P5I3 (19)

vL,d[ a 04'3.. (20)

Thus, from Equations (l), (3) and (19):

. . . . . . . . . . . . . . . . . . . . . . . . . . . . .

At constant power draw, the flow generated by the stirrer is proportional to the stirrer diameter to the power 413. In other words, the model shows that larger stirrers are more effective for foam disruption.

The effect of the different media is largely in gas hold-up and equilibrium foam height (see also below). Interestingly, the range in which vL dl influences the foam height is practically independent of the medium for a given stirrer (see Table 5) . The correlations according to Equation (17) for all media with the 120 mm Rushton turbine are graphically presented in Figure 8. The similarity of these correlations suggests that the presented model is also adequate for describing the SAFD technique in different media, despite differences in gas hold-up.

For design purposes, a minimum value of vL,dr must be striven for in order to minimize foam. This means a small distance between upper stirrer and dispersion level and a large upper stirrer at low stirrer speed for a given power draw.

Comparison with Literature Data

Veljkovic et al. (1991) published experimental data on gas entrainment from the head space in sparged agitated vessels equipped with one Rushton turbine. Some data were used in this paper to calculate a value of the liquid velocity

1026 THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997

30 T \ t medium 2 5 t

5 10 e ii I 0 0.00 0.05 0.10 0.15 0.20 0.25 0.30

VLdl [wd Figure 8 - Graphical presentation of the correlations of Table 5 for all media tested with the 120 mm Rushton turbine as upper stirrer. * 120 mm Rushton as bottom stirrer, otherwise hollow blade as bottom stirrer.

TABLE 6 Calculation of vL,dl,s from Data on Gas Entrainment from

Veljkovic et al. (1991) for Rushton Turbines

T D hU NS '$'u vL.dl.S (m) (m) (m) (h) (rps) eshmate ( d s ) 0.2 0.0667 0.133 0.004 21 1 0.113 ._ 0.3 0.1 0.2 0.0002 7 1 0.056 0.3 0.1 0.2 0.0015 11.5 1 0.092 0.45 0.15 0.3 0.0006 6.67 1 0.08 1

TABLE 7 Gas Hold-up for 10 kg Broth Weight and vZg = 0.013 d s for the

Stirrer Configurations and Media Tested ~ ~~

&G at &c at Configuration Medium 500 rpm 800 rpm

bottom stirrer (HB) I 1 I1 I L-C I I*

0.18 0.22 0.28 0.23 0.17 0.26 0.28 0.33 0.29 0.34

> 0.20 0.19 0.20 0.20 0.22 0.15 0.22 0.15

0.20 0.27 0.36 0.29 0.23 0.3 1 0.32 0.37 0.34 0.40

> 0.22 0.23 0.25 0.25 0.27 0.23 0.25 0.19

HF I 0.22 0.25 I1 0.29 0.35

HF-HF I 0.24 0.28 L-C 0.18 0.23

*RT,20 as bottom stirrer.

near the dispersion surface, VL,dl,s, which marks the onset of gas entrainment under sparged conditions. Table 6 shows that the order of magnitude for vt,dl ,s is similar to vL,dr required for foam disruption with Rushton turbines, but lower (compare with Figure 5) . From the data of Veljkovic et al. (1 99 1) it can be derived that the gas entrainment was

2(

I! e .g I C a3 E r

s - 5

C

0.25 CI A

9 0.20

r a &

0.15

0.10

CI 0.25 A

4 3 0.20 c

cn I 0.15

0.10

HB + RTei

HB + RTIS \

I

HB

8 9 10 11 12 13 14 broth mass Fg]

Figure 9 - Gas hold-up as a function of broth mass, stirrer speed and superficial gas velocity for the hollow blade stirrer (HB) as bottom stirrer and for the hollow blade in combination with one 95 mm Rushton turbine as upper stirrer (HB + RT95). Foam height as a function of broth mass, stirrer speed and superficial gas velocity for this combination. Duplicate experiments with medium I . For symbols see Figure 3.

high at N = 2Ns, i.e. vL,dl= 2vL,dl,, for constant gas hold-up. This doubling results in values of vL,d/ which compare well with those calculated for Rushton turbines in this paper. Note, that the vessel geometry used by Veljkovic et al. (1991) was quite different from the geometries used here (compare Tables 1 and 6).

GAS HOLD-UP

The target of SAFD is foam disruption through stimng, thereby enlarging the working volume of the bioreactor. However, faster stirring results in a higher gas hold-up (Whitton and Nienow, 1993). If gas a n d o r foam entrainment occurs, this would lead to an even higher gas hold-up (Veljkovic et al., 1991). But a high gas hold-up results in a low working volume, i.e. maximum broth mass, of the bioreactor. Therefore, the effects of stirring and stirrer configuration on gas hold-up were evaluated during the SAFD studies.

THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997 I027

Increasing the stirrer speed resulted in an increase in gas hold-up (see Table 7). When the hollow blade was used as the only (bottom) stirrer in Medium I, the gas hold-up was lowest (see Table 7). Adding more stirrers gave an increase in gas hold-up, which may not be explained by the increased power input alone (see below). However, this increase was disproportionally large for low broth mass, i.e. in the range of broth mass in which SAFD works (see Figure 9). For all multiple stirrer configurations, the disproportional increase in gas hold-up occurred with decreasing broth mass, i.e. with decreasing distance between dispersion level and upper stirrer. This finding is consistent with reports on gas entrainment (Hsu and Chang, 1995). Therefore, gas entrainment (foam entrainment) from the head space must have occurred during the experiments reported in this paper.

General correlations found in literature on gas hold-up, such as:

EG oc ,P v y . . . . . . . . . . . . . . * . . . . . . . . . . . . . . (21) % with p = 0.33 and y = 0.67 (Whitton and Nienow, 1993) did not seem to apply (data not shown). Matsumura et al. (1978) have shown that for systems with high stirrer speeds and with gas entrainment, quite different values for the coefficients p and y in the correlation should be used and that vsg in Equation (21) should be corrected for the entrained gas. Another reason for the lack of fit of Equation (21) might be that the estimated and not the measured specific power draw was used in the correlations. For example, Hsu and Chang (1995) found that the power draw for pitched blade stirrers pumping downward drops dramatically when gas entrainment occurs. In general, the biotransformation medium showed a 30 to

40% lower gas hold-up than the artificial media (see Table 7). On the other hand, the equilibrium foam height was 2 to 3 times higher in the biotransfomation as compared to the artificial media. Therefore, fiuther work on developing artificial foaming media should be conducted.

Conclusions

Through stirring, reduction of the height of the foam layer on low viscous broths can be achieved. This has been demonstrated using an actual biotransformation process and artificial foaming media. A simple mechanistic model for calculating the liquid velocity near the dispersion surface related to stirring and correlating this velocity to the foam height has been presented in this study. This model for stirring as a foam disruption, SAFD, technique does not account for the complex flow patterns which are stirrer type dependent. Consequently, the correlation between liquid velocity and foam height is stirrer type dependent. But for Rushton turbines with different diameters similar correlations, more or less independent of the medium were obtained, suggest- ing the independence of geometry of the model.

The recommendations from this work for reducing the height of the foam layer on production scale are, in general: 1) Reduce the superficial gas velocity by raising the head pressure andor reducing the air flow.

2) Raise the stirrers or lower the broth mass. 3) For a given size of the stirrer motor and thus power draw, reduce the stirrer speed and increase the (upper) stirrer diameter.

Apply the SAFD technique:

4) For a given stirrer configuration, stir as fast as possible, but check the gas hold-up. 5) Retrofit the bioreactor with another stirrer configuration. From the stirrers tested here, the hydrofoil Lightnin A315 had the best foam reducing capacity per unit power draw. Although not tested, the strong parallels with gas entrainment suggest other measures such as reducing the baMe number, the baffle height or the baMe width to enhance foam entrainment (compare with Tanaka and Izumi, 1987).

Acknowledgement

The authors wish to thank Rob van der Lans from the Technical University Delft for his suggestions and critical review of the manuscript. This work was supported by grants from the Swiss Federal Office for Education and Science and was carried out for the pro- ject Bioprocess scale-up strategy based on integration of microbial physiology and fluid dynamics in the Biotechnology Research and Technological Development Programme of the European Union.

Nomenclature

A ,

D = stirrer diameter (m) Fl, g h

hf L = blade length (m) m b = broth mass (kg) N = stirrer speed (s-) N ,

P Po = power number (-) Q , = gas flow (m3/s)

fL v ~ , ~ ,

= vertical surface area of cylinder with diameter T/2 above

= gas flow number, Q,/ND3 = acceleration due to gravity (m/s2) = height from the middle of the upper stirrer to the

= equilibrium foam height (m)

upper stirrer (m2 )

dispersion surface (m)

= dimensionless pump capacity coefficient for stirrer

= power draw of a stirrer (W) discharge flow, defined by Equation (3)

= discharge flow induced by stirrer (m3/s) = diameter of bioreactor (m) = horizontal radial liquid velocity near the dispersion level

above upper stirrer at distance TI2 from axis calculated from the discharge flow of the upper stirrer (m/s)

= superficial gas velocity ( d s ) = volume of liquid (m3)

v d = dispersion volume (L) V , = bioreactor volume from bottom till top of upper stirrer (L) W = blade width (m) x = material thickness of stirrer (m) y = correlation parameter, defined by Equation ( 1 8) (m) z = correlation parameter, defined by Equation ( 18) (-)

Greek letters

)

a = correlation coefficient (m) /3

y = correlation coefficient (-) E, &,,, 6 = surface tension (Nlm) p = density (kg/m3)

Subscripts and superscripts

0 95 120 g = gassed conditions S u = ungassed conditions

= correlation coefficient ( d s ) in Equation ( 1 5) , (-) in Equation (21)

= gas hold-up calculated from ( Vd - rnh/p)/Vd (-) = specific power input (Wkg)

= minimum velocity above which value hf = 0 = diameter of stirrer: 95 mm = diameter of stirrer: 120 mm

= minimal parameter value for the onset of gas entrainment

1028 THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, 1997

Abbreviations

HB = hollow blade, type Chemineer CD-6,6 blades HF = hydrofoil Lightnin A3 15,4 blades L-C = experiments camed out during the biotransformation PBD = 6-blade 45 pitched blade stirrer, pumping downward PBU = 6-blade 45 pitched blade stirrer, pumping upward RT = Rushton turbine, 6 blades SAFD = Stirring As Foam Disruption (technique)

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Manuscript received July 10, 1996; revised manuscript received September 22, 1997; accepted for publication October 3, 1997.

THE CANADIAN JOURNAL OF CHEMICAL ENGINEERING, VOLUME 75, DECEMBER, I997 1029

Documents

The Canadian Journal of Chemical Engineering Volume 75 Issue 6 1997 [Doi 10.1002%2Fcjce.5450750604] Frans W. J. M. M. Hoeks; Carla Van Wees-Tangerman; Karel Ch. a. -- Stirring as Foam