stirred bioreactor engineering for production scale part 1

8/17/2019 Stirred Bioreactor Engineering for Production Scale Part 1

1/15January 2012

Stirred Bioreactor Engineering forProduction Scale, Low Viscosity

Aerobic Fermentations: Part 1

By: Dr. Alvin Nienow

Senior Technical Consultant The Merrick Consultancy

Merrick & Company2450 S. Peoria Street • Aurora, CO 80014-5475

Tel: 303-751-0741 • Fax: 303-751-2581www.merrick.com

TECHNICAL PAPER


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Stirred Bioreactor Engineering for Production Scale, Low Viscosity Aerobic Fermentations: Part 1.

Alv in W. Nienow Merrick Consultancy

TABLE OF CONTENTS

SECTION PAGE

INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1ENGINEERING ISSUES IN STIRRED BIOREACTORS, PART 1 . . . . . . . . . . . . . . . . . . . . . . . . . . 2

M ASS T RA NSFER OF O XYGEN IN TO THE B ROTH AND C ARBON D IOXI DE OUT . . . . . . . . . . . . . . . . . . . . . . . 3

H OLD-UP AN D F OAM F ORMATION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4

H EAT T RA NSFER . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4

SUMMARY . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5

REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 6

NOMENCLATURE . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7

TABLES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 8

FIGURES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 9

REMAINDER TO COME


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Alv in W. Nienow Merrick Consultancy 1

IntroductionThere are many examples of important bioprocesses that fall into the category of low viscosity aerobicfermentations. They include the use of genetically-modified bacteria such as Escherichia coli to give bulkchemicals such as 1,3-propanediol; and E. coli and Pichia pastoris for the production of proteins for medicalpurposes. Indeed, though biofuels such as bioethanol, biobutanol and biogas are made under anaerobicconditions, biodiesel can also be produced aerobically from genetically-modified E. coli. Other examples include

Saccharomyces cerevisiae for the manufacture of Baker’s yeast and Corynebacter glutamicum for valine andlysine production as animal feed additives.

The bioreactor (fermenter) of choice for such processes uses mechanical agitation and a typical traditionalstirring configuration based on Rushton turbines is shown in Fig 1a in diagrammatic form. Such fermenters havebeen in use for at least 50 years and are commonly up to 250 m 3 in scale or even larger for the manufacture ofmany commercially important products. Fig 1b shows a photo with some people inspecting the internals in asmaller scale industrial bioreactor. In this case, the impellers are of the wide-blade, high solidity ratio type whichis one of the type more commonly used today as a result of much research which has shown that changingthe agitator from the Rushton turbine to newer types can lead to significant improvements in the fermentationprocess.

In these series of articles, the mixing issues which have to be considered when designing or operating suchfermenters will be discussed. In particular, the reasons why modern impellers are replacing the traditional

Rushton turbine by retrofitting into extant industrial equipment or have become the impellers of choice whennew plant is built will be explained. These impeller developments have been largely ignored in chemical andbiochemical engineering textbooks and in general even in the refereed journals except for a few specialistexceptions. It is also not the intention to give detailed design details or methods of calculation. These aspects areadequately covered in standard textbooks, especially that by Van’t Riet and Tramper, Basic Bioreactor Design1.

The essence of such decisions comes from recognising that the issues can all be considered as bioprocessscale-up where information obtained generally on a much smaller scale in the laboratory or pilot plant is used toestablish the desired commercial scale operating conditions. In other words, the industrial scale bioreactor hasto provide a suitable environment for the organism to grow and produce, based on the work conducted on thissmaller scale. In general, it is not possible to mimic on the industrial scale exactly the conditions found in thesmaller scale, so inevitably scale-up is a compromise. Essentially, however, this compromise is best understoodif it is recognised that each cell is itself a mini-factory converting nutrients, generally carbon based, into the

desired product. Thus the total production rate depends on production rate per cell multiplied by the numberof cells in the fermenter times the size of the fermenter. If the overriding importance of providing the correctenvironment for the cell is recognised, scale-up, though indeed a compromise, is based on that concept ratherthan arbitrary ‘scale-up’ rules related to fluid dynamics, which essentially ignore the well-being of the cell. That isthe approach recommended in this series of articles.

The equipment used to obtain this information on the desired environment for the cell is indicated in Fig 2. Atthe two smallest laboratory scales, the shaken microwell or shake flask (Fig 2 a and b) are used. The biologicalparameters that may be determined in them are shown in Table 1. To find the optimum, many experimentsmust be undertaken to establish the correct value of each of the many parameter; and this type of equipment isideal for such a requirement. It is also useful to establish the sensitivity to variations from that optimum as suchvariations are bound to occur on the commercial scale. These variations arise because the environment in the250 m3 stirred reactor is clearly not going to be as spatially homogeneous as the small volumes found in theshaken microwell and shake flask. One of the big advances that has been made in recent years in developingbioprocesses has been the ability to both measure and control many of these parameters in each reactor evenat these small scales. The papers of Buchs and co-workers, starting with one from 2011,2 give an excellentindication of the developments in this area of allowing quantitative data for engineering purposes to be obtainedat these small scales.

As also shown in Table 1, some other ‘biologically-specific’ parameters must be determined at the larger stirredbench scale (Table 1b) as shown in Fig 2c in order to be useful. Fed-batch bioreactors are ones in whichadditional growth medium is added over time, thereby increasing the cell concentration that can be achieved (aswell as the volume of medium). Thus, though the amount of oxygen that each cell requires to function properlycan be determined at the very small scale, to get a satisfactory indication of how that oxygen will be supplied


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in the commercial plant especially under fed-batch conditions really requires experiments to be done in similarequipment. Essentially, that is linked to the need to use stirring to provide the energy for oxygen transfer withair sparged from the base (specified by the mass transfer coefficient, k

La, as set out in more detail later) rather

than by shaking and from headspace aeration respectively. The potential for the energy input from the impellerto damage the organism, often referred to as ‘shear damage’ can also be assessed in such relatively small scalestirred bioreactors. This possibility comes about because though the bench scale bioreactor is small comparedto the commercial one, even at this scale, the turbulence in the flow (which is discussed further below) thatimpacts on the cell is similar at the small scale and the large. So the size of the cell to the scale of the turbulentflow is also similar.

Having the information on the bioprocess of interest at the scales indicated in Figs 1a to 1c, good engineeringmust then be used to determine suitable conditions at the production scale. These conditions should matchsufficiently closely those shown to be optimum at the bench scale with respect to cell and product concentration.Thus, the desired temperature, pH, dissolved oxygen concentration, etc, must be achieved by the agitation andaeration system in the bioreactor. The agitator provides the energy which gives rise to the turbulent liquid motionrequired for these conditions to be achieved. In particular, the dispersion of the air must be done so that theoxygen transfer rate, OTR, is able to meet the oxygen required by the cells to function properly. The turbulentflow also homogenises the contents of the fermenter so as to maintain the various operating parameters withinthe range that ensures the cells produce a similar bioprocess performance at the commercial scale to thatobtained at the small.

To give confidence that the commercial plant is going to achieve similar results to the smaller scales, operation atthe pilot plant scale shown in Fig 2d is also often undertaken. This overall approach is bioprocess scale-up withthe cell and its local environment as the key. The aim of much current research is to establish ways of going fromas small a scale as possible to the commercial scale whilst minimising work at intermediate scales, in particular toeliminating the pilot scale stage. This approach is greatly enhanced if the smaller scales, especially in the stirredbench scale bioreactor, are conducted in such a way that important parameters such as specific power (W/kg)are similar to those that will be used on the large scale. Choosing sensible parameters in this way for the benchscale is often called ‘scale-down’. A good scale down protocol greatly eases scale up

The remainder of these articles aim to help understand the interaction between the fluid motion generated bydifferent agitators under aerated conditions with variations in speed and power input; and how they change withscale. This understanding is important for the design of new equipment; and retrofitting and solving operational

issues on that already extant. To aid this understanding it is useful to subdivide this overall task into a sub-set ofsmaller issues. These issues will now be discussed.

Engineering Issues in Stirred BioreactorsMany of the engineering issues are generic and apply to all aerobic bioprocesses (and indeed gassed chemicalreactors in general). They can be considered as ‘physical parameters’; and those most relevant to bacterialfermentations are listed in Table 2. Table 2a sets out the quantitative parameters required for design whilstTable 2b lists those parameters which aid understanding and have helped improve large scale operation anddesign. The parameters in Table 1 on the other hand are specific to the organism being grown and will usuallybe different for each case. Achieving satisfactorily the biological parameters required by the cell in Table 1 by

judicious design and appropriate selection of the parameters in Table 2 is essentially the task of the engineer,especially with respect to scale-up. The final process engineering specification of the fermenter will be the size

(diameter, T (m), height, H (m)), the impeller type(s), number and size, D (m), the power needed to be imparted tothe broth, P (W) and the aeration rate, Q

G (m3 s-1). The way these values and the parameters set out in Table 2a

are determined will now be discussed.

For the microbial fermentations being considered here, the viscosity of the growth medium with the cells in it (thefermentation broth) essentially does not go much higher than that of water. With such low viscosities, the flowin the fermenter is turbulent at the 5 L bench, i.e. Reynolds number, Re =þND2/µ > ~ 104 where þ is the brothdensity (kg m-3), µ, its viscosity (Pa s), D,, the impeller diameter (m) and N, its speed (rev s-1). In practice, Reincreases with increasing scale. However, as the flow is turbulent, the actual value of the Reynolds number doesnot matter and turbulent flow theories can be used to analyse the fluid mechanics in the bioreactors across thescales. The topics listed in Table 2 will now be considered for such flows.


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Mass Transfer of Oxygen into the Broth and Carbon Dioxide outThe transfer of oxygen from air into a fermentation broth has been used since the 1940s when ‘depthfermentations’ were first established. It is one of the most important aspects of fermenter operation becauseoxygen is only sparingly soluble in water and therefore in the medium, which largely consists of water. If thesupply of oxygen into the broth ceases, its concentration in the broth would generally fall below the desiredvalue in less than a minute. Thus, the overall oxygen demand of the cells throughout the batch or fed-batch

fermentation must continually be met by the oxygen transfer rate, OTR (mol O2 m-3

s-1

); and the demand increasesas long as the number of cells is increasing. Thus, a maximum oxygen transfer rate must be achievable and thisdepends on the mass transfer coefficient, k

L a (s-1, though units of min-1 are also often used), and the driving force

for mass transfer, ∆CL (mol O

2 m-3). Thus,

OTR = k L a. ∆C

L 1

For oxygen transfer, the driving force, ∆CL, conceptually is the difference between the oxygen concentration

in the liquid film around the air bubbles (the concentration in equilibrium with the partial pressure of oxygen inthe bubble) and that in the broth. The latter concentration must always be held above a critical dO

2 value as

determined at the well-mixed bench scale. This concentration is usually expressed as a % of saturation withrespect to air as measured by a dissolved oxygen probe (% dO

2). Though the critical value is often less than 10%

dO2 (and sometimes close to zero), the set value is often as high as 40% dO

2. This higher value is to ensure that

dO2 never falls below the critical value throughout the fermenter since spatial homogeneity is difficult to achieve at

the commercial scale as discussed qualitatively above and quantitatively later. In practice, more oxygen transfercan be achieved by increasing the driving force by raising the O

2 partial pressure in the incoming gas stream

either by imparting a back-pressure on the bioreactor or by adding extra oxygen to the sparged air, preferably asa separate flow.

Roughly, for every mole of O2 taken up by a cell, 1 mole of carbon dioxide, CO

2, is produced. This ratio of moles

CO2 produced to O

2 consumed is called the respiratory quotient, RQ, and as suggested here it is generally ~ 1.

Because CO2 is very soluble in the broth, especially compared to oxygen, it dissolves. The value of k

La is similar

for both O2 transfer from air to the broth and CO

2 from it; but the high solubility, CO

2 makes it much more difficult

to strip out. However, stripping of dissolved CO2 is very important because above a certain value of dissolved

CO2 (pCO

2 > ~ 100 mbar), a reduction in fermentation rate or productivity is generally observed.

In addition to the driving force, the other parameter that the engineer can manipulate to control the rate of mass

transfer is the k L a. In low viscosity systems, kLa is only dependent on two parameters. These are, firstly, thepower input, P (W), into the fermenter, mostly from the impeller, and the air flow rate, QG (m3 s-1). It has been

shown that if the power input from the impeller when air is being sparged, P g (W) is normalised in terms of the kg

of broth, P g/M (W kg-1) in the fermenter, k

L a can be correlated with this parameter at different scales. The specific

power input into the fermenter from stirring is numerically equivalent to the mean specific energy dissipation rate,(ε

T ) g

(W/kg); and because this is a fundamental parameter in the modern understanding of turbulence, it will beused here in the rest of these articles. Thus, P

g/M / (ε

T ) g. Generally, for economic and biological reasons for

these types of fermentation, the specific power (mean specific energy dissipation rate) from the impeller is about1 to 5 W kg-1.

The airflow rate impacts in the same way across the scales if the superficial air velocity, v S (m s-1) through the

fermenter is used, where v S = Q

G/A where A

T (m2) is the cross-sectional area of the fermenter and A

T = πT 2/4 where

T (m) is the diameter of the fermenter. However, the amount of broth and the number of cells in the fermenter

increases in proportion with its volume. As a result so does the volume of oxygen required and the amount ofcarbon dioxide produced. Therefore in order to satisfy the mass balance for O2 transfer in and CO

2 out, the

volumetric flow rate of air needs to be kept essentially constant. If this flow rate is expressed on a per minutebasis, a typical practical value is about 1 vvm (where vvm is the volumetric flow rate of air at standard conditionsin m3 min-1 per m3 broth in the bioreactor). Thus,

v S = (vvm/60)(volume of broth, m3 )/(X-sectional area of the bioreactor, m2 ) 2

The combination of (εT ) g and v

S selected must together be sufficient to produce the necessary k

La where

kLa = A(ε

T ) g(v

S )b 3

k L a is difficult to determine experimentally especially on the large scale but within the accuracy achievable, this

equation is found to be independent of the number of impellers and their type and also of scale 3; and a and b

a


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are usually about 0.5 ± 0.1 for low viscosity broths. On the other hand, the numerical value of A (which is NOTdimensionless) is extremely sensitive to composition. Thus, a typical value of k

L a would be about 0.1 to 0.2 s-1 in

water with the addition of antifoam lowering it by a factor of up to 2; and salts increasing it up to x4 for the samevalues of (ε

T ) g and v

S. It is important to point out at this stage that though Equ 3 does not depend on impeller

type, the value of (εT ) g and v

S that can be eff iciently utilised does. Thus, impeller choice is very important if the

mass transfer requirements are to be met and this aspect will be discussed in the next article.

As pointed out earlier, the value of k L a is similar for both O2 transfer in and CO2 transfer out. Thus, providedscale-up is undertaken at constant vvm (or close to it), the driving force for transfer of O2 and of CO

2 will remain

essentially the same across the scales. As a result, the total volumetric gas flow rate into the bioreactor shouldalso be able to strip out the CO

2 to give the same partial pressure of CO

2 in the exit gas and pCO

2 in the broth on

scale-up as on the small scale, thus preventing potential problems with high values of this parameter. In addition,as the air volumetric flow rate scales with fermenter volume, at constant vvm, v

S increases with the linear scale-up

ratio (approximately with T commercial scale

/T bench scale

), thereby enhancing k L a if constant (ε

T ) g scale-up is also used.

Hold –up and Foam FormationThere is a down side to the higher superficial velocity on scale-up. These higher superficial gas velocitiesincrease hold-up (hold-up is essentially the proportion of the fermenter taken up by gas bubbles of different O

2

concentration). Higher hold-up means loss of fermenter capacity as cells are only producing in the broth and notin the gas phase. Even more problematic is that the high v

S increases the tendency to form a stable foam, which

further lowers productivity and if not controlled, may lead to broth being driven out of the bioreactor into the exitpipe, in extreme cases causing shut down. The usual way of handling such problems is to use one of a varietyof anti-foams4. Their use has two disadvantages. They are expensive and, as mentioned above, they lower k

L a.

It has been shown that the certain modern impellers of the type discussed later reduce the foaming tendencycompared to Rushton turbines5. Though not so well documented, experience at the industrial scale suggests thatretrofitting to these other impellers has increased k

L a by about 20 to 30% at the same (ε

T ) g and v

S because less

antifoam has been used. If retrofitting is carried out, monitoring the use of antifoam is a useful way of establishingthe change in running costs and also perhaps explaining the higher k

L a that may well be found in practice.

Heat Transfer Accurate temperature control is very important in fermentation processes as cells are very sensitive to thatparameter which should normally be held between about 35 to 40°C. The oxygen uptake rate largely determines

the metabolic heat release QH (W kg-1

) so thatQ

H ≈ 4.6 x 102 OUR 4

This cooling load has to be removed by heat transfer at an equivalent rate given by

QH.M = U A

H ∆u 5

where U is the overall heat transfer coefficient (typically about 2000 to 3000 W m-2 °C-1), ∆u (°C) is the differencebetween the temperature of the cooling water and the broth temperature (about 35 to 40°C) and A

H (m2) is

the heat transfer area available. U is hardly affected by the agitation conditions though larger D/T impellersassociated with lower power number, Po, impellers (as discussed later) maximise the inside heat transfercoefficient for a given (ε

T ) g.

Overall, at the commercial scale, heat transfer is often a problem as the cooling load scales with the volumeof the reactor, i.e., approximately reactor diameter T 3 whilst cooling surface area scales with T 2. Thus, if only

a cooling jacket as shown in Fig 3a is used on the larger scale, the area/volume goes down compared to thebench scale so that cooling coils are often required and sometimes cooling baffles (Fig 3b). Inability to meet thecooling requirements at the large scale is a very serious problem because it is extremely expensive to resolve andcannot be achieved by increasing the agitation intensity as the overall heat transfer coefficient, U is insensitiveto it. Thus, if it is necessary to increase Q

H because it is insufficient to meet the cooling requirements of the

fermentation, either AH or ∆u must be increased. The former requires the whole fermenter construction to be

modified and the latter needs refrigeration of the cooling water. It is clearly better to design with a high safetymargin with respect to the area in the first place. In a particular industrial example with which I was involved, therate of oxygen transfer was increased to meet the oxygen demand of the organism by introducing a slow oxygenfeed through a separate sparger (which incidentally is the best way of using oxygen to give an enhanced drivingforce). That approach was very successful as a way of enhancing oxygen transfer but the heat release was such


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that the temperature could no longer be controlled. As a result, the nutrient feed to the fed-batch fermentationhad to be slowed down below the maximum rate now achievable with the higher rate of oxygen transfer.

SummaryIn the Introduction, the need to consider the cell as the productive source in a bioreactor is introduced. In Part1, those engineering parameters which have to be satisfied if cells are to grow satisfactorily at the large scale

are discussed. In particular, they are the provision of oxygen in the broth at an appropriate concentration; andthe need to strip out to a low concentration, the carbon dioxide that is produced at a rate governed by the rate atwhich oxygen is consumed by the growing cells. The critical parameters in these two processes are the specificpower input from the stirrers and the airflow rate, of which increases in both increase the mass transfer coefficientand of the latter, the driving force for mass transfer of O

2 and CO

2. The air flow rate also impacts on the amount o

gas held up in the bioreactor and the tendency for foam to form, both of which can reduce its productive capacityFinally, the heat evolved at a rate determined by the rate the organism takes up oxygen needs to be removed inorder for the bioreactor to be cooled in order to operate at the desired temperature is discussed. The need to berather conservative with the area available for cooling at the large scale is emphasized.

In Part 2, the modern impellers that have been introduced into industry in recent years will be discussed. Inparticular, their impact on two aspects will be emphasized. The first topic will show how improved mass transfercan be achieved by the correct choice of impellers, thereby obtaining a higher volumetric productivity from a

bioreactor. The second topic will show that there is an inevitable temporal and spatial increase in the range oftemperatures and concentrations of nutrients, dissolved oxygen, pH, etc., experienced by the cells on the largerscale even if the mean value is closely controlled to coincide with that determined at the small scale. However, bya suitable feed strategy and choice of impellers, it will be seen that these variations can be greatly reduced, againgiving an improved performance.


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References1. Van’t Riet, K and Tramper, J, “Basic Bioreactor Design”, Marcel Dekker, Inc., New York, USA, 1991.

2. Huber, R, Roth, S, Rahmen, N and Büchs, J. Utilizing high-throughput experimentation to enhance specificproductivity of an E.coli T7 expression system by phosphate limitation. BMC Biotechnology, 11, (2011), 22-33

3. Nienow, AW, “Scale-Up, Stirred Tank Reactors”. In: Encyclopedia of Industrial Biotechnology, John Wiley &Sons, Inc., Hoboken, NJ, USA, DOI: 10.1002/9780470054581.eib535: Vol. 7 (2010) 4328 - 4341.

4. Nienow, AW, “Aeration-Biotechnology”, In: Kirk Othmer Encyclopedia of Chemical Technology, 5th Edition,Wiley, New York, USA, 2003.

5. Denkov, ND, Mechanisms of Foam Destruction by Oil-Based Antifoams, Langmuir, 20, (2004), 9463–9505.

6. Boon, LA, Hoeks, FWJMM, van der Lans, RGJM, Bujalski, W, Wolff, MO and Nienow, AW, Comparing aRange of Impellers for “Stirring as Foam Disruption” (SAFD), Biochem. Eng. J., 10, (2002), 183-195.


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Nomenclature A

H , heat transfer area

AT , cross-sectional area of the fermenter

AR, aspect ratio, H/T

B, baffle width

C, bottom impeller clearance above vessel base

D, agitator diameter

H, bioreactor fill level

k L a, specific mass transfer coefficient

M, mass of broth in the fermenter

N, agitator speed

OTR, oxygen transfer rate from the gas phase

OUR, oxygen uptake rate by the cells

P, power

QG , volumetric gas flow rate

QH , metabolic heat evolution rate

Re, Reynolds number (= þND2

/µ)RQ, respiratory quotient

T, bioreactor diameter

U, overall heat transfer coefficient

v S , superficial gas velocity

vvm, specific volumetric air flow rate

Greek Lettersa, b, exponents

∆CL, driving force

∆C, spacing between impellers with dual or more impellers

∆u , temperature driving force

εT , local specific energy dissipation rate

εT , mean specific energy dissipation rate

µ, viscosity

v , kinematic viscosity

þ, liquid density

u m

, mixing time

Subscriptsg when air is sparged


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Tables

Table 1 Bioprocess Specific Data (Nienow, 2010)

a) That Obtainable in Shaken Microtiter Plates or Shake Flasks

1 Media design

2 Metabolite concentrations including product inhibition*3 Feeding algorithm for fed batch*

4 Choice of pH control agents and sensitivity to pH*

5 Temperature sensitivity*

6 Growth and production patterns *

7 Oxygen-demand, CO2 production and RQ profile *

8 Heat-release rate (probably estimated well enough from OUR) *

9 Substrate utilization efficiencies *

10 Cell and product concentrations *

11 dO2, pCO

2 and osmolality tolerance*

* The main problem is the measurement and control of pH, dO2, pCO

2 and cell mass at these scales.

b) That Obtainable from Batch/Fed-Batch Bench Scale Stirred Bioreactors12 k

L a profile ( including impact of antifoam)

13 Foaming/hold-up characteristics

14 Sensitivity to fluid dynamic generated stresses from agitation or bubbling aeration

15 Stresses associated with spatial broth inhomogeneity on scale-up

Table 2 Generic Physical Parameters Required for Design/Scale-up (Nienow, 2010)

a) Parameters most essential for scale-up/design

1 Adequate rate of mass transfer (O2 in, CO

2 out)*

2 Bubble hold-up*

3 Satisfactory heat transfer for temperature control*4 Impeller power number, Po, and unaerated motor power draw, P = PoþN3D5

5 Mean specific energy dissipation rate, εT = P/M W kg-1

6 Reduced power draw on aeration, Pg ((ε

T)g = P

g/ M)

7 Good air dispersion

8 Effective bulk fluid mixing

* These values will be specific to each bioprocess

b) Parameters which aid understanding

9 Flow close to the agitator – single - and two phase air-liquid

10 Spatial variation in local specific energy dissipation rates, εT W kg-1

11 Gas phase mixing


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Figures

Figure 1a. Schematic representation of multiple Rushton impellers in fermenters (D/T =1/3; B/T =1/10; C/T = 1/4; H/T = ~ 3; 4 baffles)


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Fig 1b A photo taken during the installation of wide blade bydrofoil impellers in a fermenter.


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a) Shaken micro-titer plate

b) Multiple shake flasks


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c) Bench scale fermenter

d) Pilot scale fermenter

Fig 2 Examples of the range of scales used to establish the required commercial operating conditions.


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a)

b)

Fig 3 Cooling surface area provided by a) a jacket; b) coils

Cooling

Jacket

Hollow blade

impeller

Wide blade

hydrofoil

Cooling coil

Baffles which

may also be

used forcooling

stirred bioreactor engineering for production scale part 1

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