-From Biotechnology: The Science and the Business, Moses, V. and Cape, R. E., editors,

Harwood Academic Publishers (New York), 1991.

Chapter 13

Fermentation Technology, Bioprocessing, Scale-up and Manufacture ,

Yusuf Chisri and Murray Moo-Young

The commercialization of biotechnology-based processes for the improyement of human life would be impossible without the supporting engineering disciplines. Chemical or process engineering taking into accOunt the specificities of biological systems has developed ;nto biochemical engineering which is already a rapidly growing branch of knowledge. Biochemical

leering has a crucial role in economically transforming tfi1: labornory disco\'eries in biotechnology into large scale manufacturing. Process development time - the time between initial process conception and full scale manufacture of p-0duct for sale - may be considerably shortened by the earliest invoh-ement of engineers in biotechnology research. The type and the quality of research data, for example, can lead to significant savings in resources in subsequent stages of biorr"c.. s~ development.

This chapter is an overview of biochemical engineering base of biotechnology. The processing considerations common to many biological systems are examined. The bioreactors used in the production of biochemicals and biocatalysts (enzymes, microorganisms, cell cultures) and the fundamentals of design of these reactors and supporting srstems are treated. 1"' 'iuct harvesting, purification and other downstream pro-b.lOg operations are discussed with emphasis on newer

p'elcpments. The associated process control technologies a:" considered. Finally, an overall process dimension is pro­v Jed by ~x:lmining a complete bioprocess.

BIOPROCESS: GENERAL ASPECTS

Any large-scale operation involving the transformation of some raw material (biological or non-biological) into some product by means of microorganisms, animal or plant cell cultures, or by materials (e.g. enzymes, organelles) deri\Oed from them, may be termed a "bioprocess". The "product" of such processes may be saleable (e.g. insulin, penicillin, SCP, enzymes) or they may have little commercial value (waste tre:ltment). A bioprocess is typically made up of three steps shown i.n Figure 1. The raw material or feedstock (see Chapter 14) must be converted to 3 form which is suitable for p~ocessing. This is done in a pretreatment step which may Involve one or more of the operations shown in

Figure 1. Frequently, the well established chemic31 engineering operations suffice for the pretreatment stage and these will not be discussed further.

The pretreatment step is followed by one or more bio­reaction stages where the

PROCESS STAGES -

Raw Material

I PRETREATMENT

l

I I

I l

BIOREACTION I I

I DOWNSTREAM PROCESSING

I I

8 It

desired biotransformation takes

I

OPERATIONS

Sorting Sieving Comminution Hydrolysis Medium formulation $Ieri lization

Biomass production Metabolite biosynthesis Immobilized enzyme and

cell catalysis Bio trans format ions

Fi lira tion Centri fugation Sedimentation Flocculation Cell Disruption Extraction Ultrafi Itration Preci pi to t ion Crystalization Chromatography Evaporation Drying Packaging

Figure 1. Bioprocess stages and the commonly used operations in them.

I

Bingham

Casson

C\J

E TO Oilotanl

Z n > I Newtonian n = Il-

Pseudo-plastic n <: I./ I

;' ~ I..--./ Islope = fLo p I

I . I Yo )'1 I t

0 0 y (5- 1 )

Figure 2. Shear strcss \"s. shcar ratc plots for varIOUS flow behaviours.

may be used. For mrcdial broths the power law model is oftCl1 s:llisfactot·y. The parameters K and narc dependcnt on solids concentration; Il generally declines \\·ith the concen­

! tration of solids, i.e., the suspension bccomes increasingly : __ shear thinning, while K increases quite strongly with the . ~ solids content of the mycelial slurry. At present, however, no

satisfactory quantitati\'e relationships exist to describe the influence of mycelial solids on K and n of these fluids. For example, for broths of Aspergil!lls niger the dependence of K on solids concentration (X) according to two different correlations is shown in Figure 3. The agreement between these equations is not at all satishctory. As a result, eYen though a flow model may fit the experimental d:lta, the parameters obtained ma;' ha\'c little yalue for bioreactor design, Furthermore, all the foregoing flow models and the parameters obtained on their basis apply strictly to laminar flow. Could such K and n have any real value for process design purposes under conditions of turbulent flow? This question is particu­larly important in man;' biological slurries in which the parameters K and n are not really constant but depend on the shear rate range used in their determination. Questions such as this remain to be ans\\'ered. Alternative approaches which relate the design propertics such as oxygen transfer, mixing and heat transfer directl;' to biomass content of a slurry are useful as discussed later.

Permenter technology

10..-------------,---1

AspergillUS niger

K = 0.03 X2.48

based on Reuss ela!. (1982)

c (J)

0 1.00...

~

K= 4.3x IO-4X3.3

based on Allen (1987)

0.1 L-_..1----'---..L--L-.--lL--L...L-l~-----'

I 2

Figure 3. Dcpendence of consistency index (K) on solids concen­tration (X) in broths of Aspergillus niger.

PRETREATMENT OPERATIONS

Media Sterilization

The successful operation of most productive fermentations depends on the maintenance in culture of a single microbial species with \\'hich the fermenter was initially inoculated. To prevent contamination by other organisms the gaseous and liquid feeds to the fermenter have to be sterilized, Either phrsical elimination, for example by filtration, or destruction of microorganisms mar be used to achieve sterile feed streams.

\,\'hile ionizing radiation and chemical sterilents rna:' be used, sterilization by filtration and heat treatment are b;' far the most common techniques. Liquids such as water and salt solutions can be inexpensively sterilized by passing them through absolute filters the pore size of which are smaller than the dimensions of any contaminating particles. This method is practicable only when the liquid is free of other suspended mat~rial and its initial microbial contamination is

4 6 8 10

X (gL- 1 )

Biotechnology/The Science~nd the Business

pbcl.? The tr:lI1sform:uion may involve the conversion of a substrate to biomass or biomass and some biochemical or enzyme. Alternatively, the conversion may usc dead whole cells (immobilized or in suspension) or enzymes as the bio­catalytic agl.?ncy. Bioreactors form the core of the bioreaction stq.1. The material produced in the bioreactors must usually be processed further in the downstream section ofthe process to convert it to a useful form. Downstream processing con­sists of predominantly physical operations which arc aimed at concentration and purification of the product. The purified product may have to be in different physical forms (liquid,

" slurry, powder, crystalline) for different applications. The properties of biological materials impose significant

unique constraints on the bioreaction and dO""nstream pro­cessing stages. These stages arc treated more thoroughly later on in this chapter after a consideration of the characteristics of the biomaterials themselves.

--~ PROPERTIES OF BIOLOGICAL MATERIALS

The design of a bioprocess and the engineering of the process equipment requires a careful consideration of the physical and chemical properties of the material being handled and a delimitation of the maximum processing stresses (temperature, pH. shear forces, contamination, pressure) that the material nu~' safely tob·ate. Typically a bioprocess must operate within the physiological ranges of pH and temperature (pH - 7, tempaature ::; 37°C), the specific conditions being very process-dependent. The pressures which would normally be encountered (::; 2 MPa) in bioreactors do not seem to damage microorganisms or enzymes but carbon dioxide associ:lted toxicity effects due to increJsed solubility of CO2 at higher pressures may become important.

The sensitivity to sheJring forces mJy \"::Ir}" widely. In general, bacteria and yeasts which grow JS small individual cells arc quite shear tolerant. GeneticJlly engineered species and wall-less mutJnts arc frequently susceptible to shear damage due to their weaker cell wall/membrane structures. Filamentous bJcteriJ and mycelial fungi which have larger particle dimensions do show signs of mechanical damage in high shear fields. Similarly, plant and mammJliJn cells are more sensitive to shear. Enzymes, with the exception of multienzyme complexes and membrane associated enzymes, Jre. not damaged by shear in the Jbsence of gas-liquid inre~f~ccs. "'B~use nearly all processing operations must handle liquids

and slurries a more in-depth tre:1tment of biofluids follo",·s.

Biological Fluids

Biofluids and slurries hll into two categories: (i) Ne",·toniJn fluids such as water, honey Jnd most bacterial and rcast

fermentJtion broths; and (ii) non-Newtonian media such as polysaccharide fermentations and the broths of Streptomyces, Aspergilli and Penicillia.

Newtonian fluids At constant temperature Jnd pressure, Newtonian fluids have J constJnt viscosity irrespective of the shear rate. for these fluids shear stress (t) and shear rate arc linearly related. In laminar flow

t = ~L Y (1)

Plots of T vs. yare straight lines of slope equal to the viscosity, ~lL'

Non-Newtonian fluids Viscosity of these fl uids is dependent on the rJte of sheJr. This dependence is commonly described by the power law "­model:

t = KyO ":-., ---''';~'

where K and n Jre the consistency index and the flow beHav­iour index, respectively, of the fluid. By anJlogy with eq. (1) In Jpparent viscosity (~l;lp) CJn be defined for the non~

Newtoni:lI1 fluids:

(3)

for n = 1, eq. (3) reduces to constant viscosity form and the fl uid is Newtonian. for n > 1, the fluid becomes increasingly viscous with shear and it is termed dilatant or shear thickening.. When n < 1 the fluid is shear thinning or pse/fdoplastic. . Many biological media display pseudoplastic behaviour.

The shear stress vs. shear rate plots for the various flo"'­behaviours are shown in Figure 2. The slope of a line joining any point on these plots to the origin is the JppJrent viscosity. Clearly, the Jpparent viscosity increases (dilatJnt) or dcc; '5

(pseudoplastic) with shear rate for non-Newtonian fluiq J Certain fluids do not flow until the applied shear s' A's

exceeds a minimum VJlue (To) known as the "yield 'Str~" This t)Tpe of behaviour may be described bv the Bing/;'s . , . ,vn plastic flow model or by the Casson equJtion:

T = To + K )T (Bingham plastic) (of} C

y'T = y:r;, + Kc yy (Casson model) (5)

where Kc is a constJnt kno",-n as CJsson viscosity.

Estimation of viscosity For suspensions of yeJsts Jnd bacteria gro""ing JS individuJI cells suspended in a wJter-like medium, equations of Einstein

~lS = ~lL (1 + 2.5 ES)

and of Vand

(7)

168

wm'~;u?'BFW'f?rWZW:mfC'S;m7

sterilization is calcubtcd by gnphical integration of kJ VS.

time profiles

-In - J kd dt + kd thold + (13 ) No TOlal(N)

To

In continuous sterilization very rapid heating and cooling is obtained and the contribu tions of heating and cooling periods to sterilization is generally disregarded.

Heating Hold I Cooling' w n::

T s --------­I I

::> I I r- I I <t n::

I I

I I

w Q.

I I

I I

~ I I uJ r- To

I ... t heat

I I I., .....

t hold I

I t cool ~I"'"

i -'1

to TIME

Figure 5. Time-temperature profile.

Continuous sterilization

Ad\'antages associated ""ith HTST sterilization, rapid heating and cooling and precise control of holding time makes con­tinuous sterilization the preferred method whenever it can be employed. The kinetics of sterilization are identical to the treatment gi\'en in the previous section. The raw medium is 'eated to the sterilization temperature either by continuous

:- ~team injection or by a high efficiency heat exchanger (plate or spiral exchangers). This is followed by a holJing coil where the sterilization temperature is maintained for a time equalling the residence time of the coil. Either flash cooling or indirect heat exchange returns the feed to the fermentation temperature. The steam injection and indirect heating schemes are illustrated in Figure 6.

For a continuous fermentation feed flow of Q m\-l, N may be easily calculated for a given level of contamination, e.g., one contamination during operation time top, as

N = lIQ top' (14 )

Hence, the holding time necessary for a given NINo, and the length of holding coil rna)' be determined.

The design of the holding coil requires careful attention, The velocity of flow is not uniform across the cross section of a pipe and the flow is always faster at the axis of a straight pipe. As a result, the residence times of different clements of fluid in the pipe can be diffaent. A most conservative estimate of the residence time is obtained from the equation

Fermenter technology

L (15)t r =

Un1 .\X

""here L is the length of the holding coil and U m .1X the centreline velocity of Ho"" in the pipe. The maximum velocity U m .1X

depends on the Reynolds number of flow

PL U d Re = :......::--- (16) ~lL

""here U is the average flo",' velocity through the pipe. \\'ell­de\'e!oped turbulent flo"" (Re » 2300) is desired in th<' pipe to minimize the difference between U and UOl.lx' Values of U fUm." as a function of Reynolds number are gi\'en in Fioure 7 for straioht, circular pipes. In helical coils laminar

~ ~ .

flo"" persists to significantly higher values of Reynolds number than in straight pipes.

Air Sterilization Aerobic fermentalions require a continuous supply of large quantities of air or oxygen. The gas entering the fermentc'r must be free of contaminants such as bacteria, spores, bac­teriophages and other microorganisms. Similarly, the exhaust gas from a fermenter must be treated to remo\'e micro­organisms and spray particles which can be potentially harmful to plant personnel and the environment.

Sterilization of fermentation inlet and exhaust gas is achiewd predominantly by filtration on which depends the fail ure or success of a fermentation operation. The smallest particles which need to be remO\'ed from the air are bacteria and viruses. The smallest bacteria are - 0.1 f.lm in diameter; viruses are typically less than 0.3 ~lm and may be as small as 0.0-1 ~lm. Either depth filtration or absolute filtration may be used to free the air of umvanted particulates. Depth filtration depends on passing the gas through a bed of packing such as compressed glass wool or other fibrous material. The spaces between fibres are larger than the dimensions of the particles to be remowd and particles penetrate the filter bed to various depths. The total filter depth is such that the required reman] is achie\'ed. Absolute filters, on the other hand, have openings which are smaller in size than the dimensions of the smallest particle to be retained. Particle remonl is by a sieving action. Porous polymer membranes, ceramic and metal membranes are used as absolute filters. Polymer membranes in the very narrow pore size distribution can be produced by subjecting non-porous polymer films to bombardment by high energ~'

nuclear panicles. Howe\'er, absolute membrane filters are easily fouled and produce relatively high pressure drops. Prefilters are used prior to the air filters to remove gross contamination such a dirt, oil and water droplets and foam to

extend the operational life of the filter.

Depth filtration

Sewral different mechanisms of particle retention operate in a compacted bed of fibrous material: direct interception of

171

~iotechnology/The Science and the Business

ow. For some liquids such as blood serum filter sterilization nay be the only viable technique if denaturation of its highly abile constituents is to be prevented. On the other hand, :oncentrated solutions of sugars and highly contaminated nedia (e.g. molasses, cornsteep liquor) are commonly heat terilized. Thermal denaturation of contaminating organisms, \·hen properly carried out, is among the most effective nethods of sterilization.

:-feat Sterilization

fhermal denaturation of one or more enzymes in the con­aminating microorganisms is used to render them non-viable. fhe feed is brought to a sufficiently high temperature and leld there for a certain time to ensure the destruction of the nost resistant contaminant. The process may be conducted )atchwise or continuously.

rJatch sterilization

fhe rate of destruction of microorganisms follows the first­)rder kinetics:

(8)

iVhere kd. is the specific death rate and N the concentration of :he contaminating organisms at any time. The time (t) required :0 reduce the concentration of contaminants from an initial value No to some value N is obtained from the integrated :orm of eq. (8):

(9)

[ypically, No is 105 -109 /mL. The final concentration N iepends on the degree of acceptable fermentation failure ;ince at some point the cost of further reducing the risk of :ontamination would exceed the expense of a lost fermen­:ation. For a 250 m3 fermentation with an acceptable failure 'ate of one in fifty fermentations the final level of contami­lation would be one microorganism in fifty fermenter volumes

N = 1/(50 X 250) = 1/12500 m3 == 8 X 10- 11 per mL

[he thermal death rate constant kd is a function of temperature

kd = A e-t>E/RT (10)

vhe~e t..E is the activation energy for the destruction of a Jarticular microorganism, A is its Arrhenius parameter and T s the absolute temperature. Bacterial spores such as those of fJacillus stearothermophilus and Clostridium botulinum are ;ome of the most heat resistant, and sterilization processes lre designed to be effective against them. The activation :nergy for destruction of these spores is 2.5 - 2.9 X lOs ] ,mol- I and A is - 1.6 X 1036

S-I. Because kd increases with

]0

temperature, the higher the temperature the shoner the treat­ment time needed to achieve a given b·d (N/No ) of destruc­tion as illustrated in Figure 4. However, the processes ,,·hich lead to microbial inactivation also cause destruction of heat labile essential nutrients in the feed. Denaturation of nutrients (kn ) is also temperature dependent

kn = An e-t>En/RT (11)

o Z "­Z

100°C

IISOC

A = 1.6 X 1036 5

b.E =2.9 X 108 J kmol- I

0.1

0.001 =- =-'::" ----'I -'---_-..J

a 0.5 1.0 1.5

HOURS

Figure 4. Time needed to obtain a given level of destruction (N / No) at various sterilization temperatures.

From equations (10) and (11) it follo,,'s that

kd = ~ e(t>En - 6E)/RT (12)kn An

The activation energy for the thermal deactivation of most nutrients (t..En ) is substantially lower than t..E for microbial deactivation. For example t..En for thiamine hydrochloride (vitamin B6) is only 9.2 X 107

] kmol- I• Hence, in order to

maximize microbial destruction relative to nutrient loss (i.e. achieve higher kd/kn ), sterilization at high temperature for a shorter time is indicated. This is the basis for HTST or high­temperature-short-time sterilization. Because of the short exposure times (of the order of seconds), HTST is best implemented in a continuous flow mode.

Either direct heat (steam injection) or indirect heat (coils, jacket) may be used to heat the fluid to the sterilization temperature where it is held for some time (holding time) followed by cooling. The time-temperature profile of such a process is shown in Figure 5. 1':­

In batch sterilization heating and cooling times are relatively' long and some sterilization occurs during these periods. During heating and cooling kd varies with time and the total

Fermenter technology

of all viable organisms. In practice, each depth filter thickness may haw a DOr penetration of 0.001 % or less and since se\'eral 1:1\'ers are often used in series, the actual penetration is much I;,,·er. Hence, depth filters are an effective means for air sterilization. In some applications a depth filtration step mav be followed by an absolute filter to gain additional cOl~fidence on air purity. A typical packed filter arrangement in fermentation application is shown in Figure 8.

BIOREACTION

Fermentation Process, Growth and Production

A sterilized batch fermenter containing a properly developed fermentation medium (C and N" sources, micronutrients), sufficient aeration, supplies of pH control chemicals, antifoams and associated systems must be inoculated with the desired microbial species to initiate the fermentation. The inoculum consists of a microbial suspension in rapid exponential growth added at a concentration of 5-1 0% by volume to the fermenter. The slower growing the organism, the larger the volume of inoculum used to avoid having long fermentation times (costs) in the production vessel. Because industrial fermenters tend to be quite large, inoculum preparation from agar slants often requires several fermentation steps: shake flasks, seed fermenter and secondary (or tertiary) seed fermentation. In some instances quantities of spores for inoculation produced in a seed stage arc blown directly into the larger SC3le vessel with the ingoing air.

Following inoculation the growth of the microorganism follows the typical pattern illustrated in Figure 9. Inoculation of cells into the fermenter often results in a period where there is no increase in cell number: this period is known as the "lag phase". The length of the bg phase is related to the growth histOry of the inoculum, the composition of the medium and the size of the inoculum. The composition of the media used in seed and production vessel should be identical to avoid or eliminate excessive lag. Additionally, as pointed out earlier, rapidly growing cells (late exponential growth phase) should be used for inoculation and the volume of the inoculum should be such that possible osmotic st:.9-ck effeels on dilution in the brger vessel are minimaV The existence of a lag phase shou Id not be taken to mean lack of metabolic activity in the cells; in fact, the lag phase is pre­paratory to rapid exponential growth and is essentially an ad3ptation period. Ne\'enheless, the lag phase is unproductive with respect to fermentation time in the production vessel and fennentation optimization aims at reducing or bypassing the lag. The bg phase is followed by exponential growth during which the cell number (or mass) increases exponentially with time. The cell mass and cell number grO\\'th r3tes are not necessarily equal.

I Log phase l/) l/) 1 :Exponential : Stationary Death

I growth I _..:.p_h_a_s_e_--..;.._ phase o E

Q.l U

Q.l

.0 o >

I I I I I J

I I I

o FERM ENTATION TI ME

Figure-9. Typical progression of microbial growth.

Increase in cell mass (X) during exponential growth often follows the equation

dX (18)

dt

where ~l is the specific growth rate and kd the specific death rate. During exponential growth ~l » kd and eq. (18) reduces to

dX ~ = ~lX (19)

For a cell mass concentration Xo at the beginning of expo­nential growth (Xo usually equals inoculum concentration in the fermenter) and taking the time at which exponential growth commences as zero, eq. (19) can be integrated to

yield.

X In - = ~lt (20)Xo

Hence, a biomass doubling time td can be shown to be

In 2td =-~ (21 )

11

Similarly, a cell number doubling time or mean generauon time tg is giyen by

In 2 t g =-- (22)

flN

where flN is the specific cell number growth r3te. In bacteria, where cell division leads to two identical cells,

flc-; and fl "'ill be the same. For yeasts, moulds, plant cells and other organisms ~lN and fl are not always equal. The specific growth rate is characteristic of the microorganism and is a function of growth environment including temperature, pH,

173

-- -­- -- ­

Biotechnology/The Science and the Business

To vaccum

Holding coil

Heater HoldingFlash evaporator coil

sterile Steam mediumRaw

medium injectionSterile

medium

(a) Direct steam injection (b) Indirect heating

Figure 6. Stcrilization processes: (J) direct steam injcction; (b) indirect heating.

Raw medium

..E 1.0..----------------------, ~ ::J

o 0.8 ~ a:

~ 0.6 u o -l W 0.4 L----L........l-.L..J...LUJL.I.----L---l...-L...l...LU.l..l..-----J'---'--Ll-Ll..LJ.J

> Id Id 104

105

REYNOLDS NUMBER. Re (-l

igure 7. U/Umax vs. Reynolds number in straight, circular pipes.

articles by filter fibres; inertial impaction on fibres of particles I ith sufficiently high momentum that they are incapable of following the gas stream as it flows around to avoid the fibre ~nd t,hey impact on the fibre; small particles mm'e around in rhe gas by diffusion and Brownian motion, they eventuall)' ~ollide and are retained on filter fibre. Other modes of particle emoval mechanisms may occur to various degrees. Thl: emoval of particles with filter depth follows the equation

No kIn N = r z (17)

72

where No and N are, respectively, the initial and final levels of contamination in air, z is the filter depth to achieve N contamination level and kr is the filter constant. kr depends on bed void fraction, fibre diamcter, velocity of flow and temperature. formation of condensate affects filter performance and should be prevented.

The air filter is itself sterilized by direct or indirect steam heating. Chemical sterilization is also practiced. ~

A common test of filter performance is based on penetration of dioctylpathalate (DOr test) particles mainly of :::; 0.3 ~lm

size. A penetration of 0.003% or less corresponds to removal

Prefi Iler ( Humidify control)

Air 1------,

_steam

Filter one-way valve

steam

Condensate drain ..

" ... .

Fermenter

Figure 8. Air filtration using a packed filter.

-Fermenter technology

z 0 ".- , ­I- ,/

/

<I: / /cr:::

I- //

Z / Product w /

/(.)

Z 0 --- ( a)(.)

TIME

/ /

/ ".....

/ /

/ /

/ /

/ / Product

/

( b)

TIME Figure 11. Relationship between growth and product formation: (a) growth associated product formation; (b) non-growth associated production.

of a purer carbon source such as glucose compared, say, with molasses may reduce purification problems and simplify pollution control and waste treatment.

The medium must provide sufficient carbon, nitrogen, minerals and other nutrients to yield the required amount of cell mass and product. Minimum requirements arc estimated from the stoichiometry of growth and product formation. In general,

C-source + N-source + minerals + specific nutrients + O 2

(vitamins, hormones) -> cell mass + product + CO2 + H 20 (24)

Most nutrients arc supplied at levels well abo\'e the minimal needs. Other considerations relating to fermentation feed­stocks have been examined in Chapter 14.

BIOREACTORS

Fundamentals of Mass and Heat Transfer

Transport of mass and heat are encountered not only in bioreactors but also in most other processing operations. Heat sterilization of fermenters and temperature control during a ferment:1tion are both dependent on heat transfer phenomena. Similarly, the transfer of oxygen from a gas phase into a liquid and within the liquid to the biocatalytic particle are problems of mass transfer.

Gas-Liquid Mass Transfer

The transfer of oxygen from the gas-phase to the micro­organism suspended in the gas-liquid dispersion takes place along a certain pathway. The most general transport route is

12) G':'S -lI0UID INTERFACE

OXYGEN

141 BULK LIQUID

{3l LIQUID FILM

AIR BUBBLE 1

(6) CELL- LIQUID INTERFACE

(8) SITE OF BIOCHEMICAL REACTION

Fipure 12.. Oxygen transport path from the gas bubble to the mlcroorgantsm.

depicted in Figure 12 which shows that eight resistances to oxygen transfer can exist:

1. in a gas-film inside the bubble; 2. at the gas-liquid interface; 3. in a liquid film at the gas-liquid interface; 4. in the bulk liquid; 5. in a liquid film surrounding the cell; 6. at the cell-liquid interface; 7. the internal cell resistance; and 8. the resistancc at the sites of biochemical reaction.

Not aU these resistances are significant. Thus, in practice bioreactors operate at such levels of turbulence in the fluids that convective transport dominates in the b~dy of the

Fennenter technology

:.. •..~~~r" Tn other words, the mass transfer IS liquid Since the transfer rate and the flux are related by

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Biotechnology/The Science and the Business

medium compositIOn and dissolved oxygen levels. Some typical doubling times for different classes of microorganisms are given in Table 1. Exponential growth is followed by a stationary phase during which the growth and death rates are equal. Exhaustion of a growth-limiting nutrient in the medium or accumulation of toxic material are possible causes of onset of the stationary phase. Eventually the culture enters a death phase in which cell lysis or some other mechanism of loss of cell viability overtakes growth.

Table 1 Typical Doubling Times of Some Industrially Important

Classes of Microorganisms

td (minutes)

Bacteria 45 Yeasts 90 Moulds 160 Protozoa 360 Mammalian hybridoma 630-1260 Plant cells 3600-6600

Effect of temperature on growth

Depending on the optimum temperature for growth, micro­organisms are classified as psychrophiles, mesophiles and thennophiles. Typical optimum growth temperature for these is giYen in Table 2. Actually, there is a range of temperatures ncar the values given in Table 2 over which these classes of organism grow; the exact optimum growth temperature depends on the microbial species and other growth conditions. The efficiency of conversion of the carbon source to cell mass is temperature dependent and declines with temperature. Maximum growth yield is obtained at temperatures lower than those for maximum growth rate. Furthermore, tempera­ture optima for growth and product formation are not necessarily the same.

Table 2 Typical Optimum Temperature for Growth

Psychrophiles -15 Mesophiles - 37 Thermophiles -55

Substrate concentration effects on growth

The effect on growth of the concentration of a growth­limiting substrate, such as a carbon source for example, often follows the behayiour shown in Figure 10. The specific growth rate !l increases with substrate concentration until it is no

longer growth limiting. The curve in Figure 10 is described by the Monod equation

S (23)

j-Lt

~m ---------~------

~m

2

s~

Figure 10. Effect of substrate c~ncentration on specific growth rate. Ks is the substrate concentration at half llm o

where ~lm is the maximum specific growth rate and Ks is the saturation constant. Numerically, Ks is the substrate concen­tration corresponding to !lm/2. Thus, growth on a given substrate may be described by two constants: ~lm and Ks· However, at high substrate concentrations inhibition of growth due to substrate may be encountered. High substrate concentrations may adversely affect product formation in many fermentations.

Consideration of the relationship between cell growth and product formation is essential to the successful conduct of a .,----' fermentation. Two simple, extreme possibilities are growth­associated product formation (Figure 11(a)), in which case the product formation results from primary energy metab­olism, and nongrowth associated production where product concentration is proportional to the quantity of biomass but not to growth rate (Figure 11 (b)).

Growth Media

Careful formulation of growth and production medium is a prerequisite of successful fermentation. Microbial nutritional and environmental needs have to be met as well as several technico-economic constraints. Medium development aims to

maximize product yield and cDncentration at minimum medium cost. Although traditional emphasis is on the fer­mentation step, the choice of medium affects downstream and upstream (pretreatment) activities and medium design should be carried out in an overall process context. The usc

174

r

It = K ;,n-l (53)• ,If' I

while the average shear rate (y) in the stirred vessel is often approximated by the equation

;/, = k, K (54)

Substitution of equations (53) and (54) in equation (52) leads to a modified Reynolds number:

-N'-n D'PL - ,­(55)Re, = K kn-1

I

The Powernumber-Re,'nolds number behaviour for non­Newtonian media is ~imilar to that described for the Newtonian fluids. The constant k, in eg, (54) is dependent on the system geometry and to some extent on the rheology of non-Newtonian media. The dependence of k i on the flow index (n) of non-Newtonian fluids has been reported to be

4n )n~i

Fermenter technology

sterile air. These bioreactors employ multiple impellers, usually Rushwn turbines. \X'ith some modifications the design methods de,doped for the standard reacwr can be usefully employed for other geometric configurations. 1\lultiple impellers are located on a single shaft with a minimum distance of one impeller diameter between impellers. Two diiferent types of impellers may be used on the same shaft and are in fact beneficial in some cases. For example, a combination of radial flow Rushton turbines and axial flow propellers is sometimes used to provide improved mixing in vessels \\'ith HIdT > 1. Recent investigations inw other impeller designs ha,'e sho\\'n that impellers such as Prochem hydrofoils and InrerMIG provide higher mass transfer at lower power con­sumption than Rushton disc turbines. Some of these newer impeller types are depicted in Figure 18. In many cases existing stirred tank bioreactors are known w have been upgraded by changing the Rushton turbine in favour of the newer impellers. The lower shear characteristics of the latter

k= K' -~ I

( 3n + 1 (56~y a:~~ added advantage. ~-- /

where K' is impeller-tank geometry dependent. Some typical values of k j are given in Table 5.

Table 5 Values of k j (eq. 54) for Various Impellers

k1

Impeller (- )

6-Bladed turbine 11-13 Paddles 10-13 Propeller -10 Helical ribbon -30

Quite distinct from an average shear rate in the tank, a maximum shear rate on the impeller also exists, Considerations of maximum shear to which a fermentation culture may be exposed without harm are particularly important for fragile biocatalysts \\·hich also tend to have large particle sizes (animal cells, plant cells, fihmentous organisms). The maximum shear rate on a Rushton turbine in Newtonian media has recently been expressed as

y = 3.3 Nl.5(D j 2

PL )[1 (57) ~tL

which applies for Rei = 100 to 29,000.

Other Stirred Tank Configurations

The "standard" configuration of stirred tanks is not the commonly used geometry in bioreactor applications; instead tanks with height-to-diameter ratios of 3: 1 and 4: 1 are more common because they permit better utilization of the expensive

Pneumatically Agitated Bioreactors

Although traditionally in extensive use, the mechanic"ally stirred bioreactors have several significant limitations \\'hen compared to pneumatic bioreacrors which are agitated b}' gas injection. A comparison of pneumatic bioreactors (airlifts and bubble columns) with stirred vessels is provided in Table 6. Numerous advantages of the gas-agitated reactors have led to a clear preference for them particularly in the newer bio­technology applications im'olving fragile material. The following sections examine some of the design considerations for bubble column and airlift bioreactors.

Bubble Columns and Modified Bubble Columns

Bubble columns are among the simplest of gas-slurry bio­reactors. A gas-sparged pool of Ii uid with heioht-to-' er ratio well a ove unlt\- constitutes a bubble column ~gure 19). Several modifications to the basic design are possible, however, and Figure 19 illustrates some possible configurations.

The energ" input to the fluid in the bubble column arises redominantly from isothermal expansion of the injected gas

and epen s on t e super cia gas ve ocity

Pc V = PL g U G (58)

L

The main reactor performance characteristics gas holdup (f), specific gas-liquid interfacial area (ad, overall volumetric mass transfer coefficient (kLad, mixing, axial dispersion (Ed and heat transfer - are controlled by gas flow and hence also by the energy input. A wide range of gas velocities may be used; howe,'er, the maximum velocity should be less than the

181

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liquid and hence the associated resistance (i.e. resistance 4 in Figure 12) can be ignored. Similarly, for single cells or dis­persed mycelia the resistance due to the liquid film on the surface of the cell (i.e. resistance 5 in Figure 12) may be neglected. This is so in spite of the fact that the density differences between the microbial cells and the s~spending

fluid are very small and, consequently, there may be a stagnant liquid film surrounding the microbe. The minute dimensions of a microbial cell and its large surface area are the reasons for a negligible cell-liquid film resistance. The cell-liquid interfacial resistance and the resistance at the reaction sites (resistances 6 and 8, respectively, in Figure 12) can both be disregarded because of the active oxygen transport through the cell membrane as well as the rapid rates of biochemical reactions. An intracellular resistance (i.e. resistance 7 in Figure 12) can also be discounted since the enzymes for terminal respiration are located in the cell membranes rather than in the protoplasm, and again because of the small size of microbial cells. This efu!!inatesa,ll the transport resistances ~~~t!h()!!-~~rouE,d.the ~as-li9uid interface. The oxygen transp02:t.,p~2J?lernisther~by reduced to thatof the gas-h~

"Interfacial mass transfer. .

The mass transfer models

The region in the vicinity of the gas-liquid interface may be visualized as consisting of adjacent, stagnant, gas and liquid films of some finite thickness as depicted in Figure 13. According to this two-film model, the resistance to transfer in each phase is localized in the thin films close to the interface. The interface itself is assumed to offer no resistance to mass transfer; the interfacial concentrations are therefore determined by the equilibrium relationship. Mass transfer through the stagnant films is assumed to be solely by molecular diffusion and thus at steady state linear concentration profiles exist in the films (Figure 13). For this situation the mass flux (Jo,) of the diffusing species is related to the concentration gradient (~C) in the film and to the film thickness (6) in accordance with Fick's first law:

DJoz= "6 ~C (25)

where D is the molecular diffusivity of oxygen in the film. The ratio D/6 is known as the "mass transfer coefficient", k. Equation (25) may be written for each of the two films.

Joz= kG (CG - C Gi) (26)

= kL (CLi - Cd (27)

where kG and kL are the gas and the liquid mass transfer coefficients, respectively. Since the interfacial concentrations are in equilibrium, the flux may be expressed in terms of the overall concentration driving force as follows:

(28)

DIRECTION OF DIFFUSION .. INTERFACE

BULK BULK GAS GAS LIQUID LIQUID

FILM FILM ':J !

CG I "STAGNANT" I FILMS I I I I

II T

I I I I CLI

II 1--8G 8L --J

Figure 13. Oxygen concentration profile in the gas-liquid inter­facial region.

where KL is the O\"erall mass transfer coefficient based on liquid film. C:' is the equilibrium concentration in the liquid, which, for a sparingly soluble gas such as oxygen, is related to C G by the equilibrium relationship known as Henry's law:

CG = He:' (29)

where H is the Henry's law constant. From equations (26), (27), and (28), and the knowledge '-..-/

that CLi = H CGi, it can be shown that

1 1 1 - = - + ~ (30)KL kL HkG

For sparingly soluble gases, such as oxygen, H is very much larger than unity. Moreoycr, kG is typically considerably larger than kL because the gas phase diffusivities are vastly greater than those in the liquids (d. Doxrgon/air = 10"' Doxrg~n!walcr at 20°C), and at the same time, the gas-pha~e film thicknesses are smaller than those of the liquid films. Under these circumstances the second term on the right hand side of equation (30) becomes negligible and the equation reduces to

1 1 (31 )K = k

L L

This implies that essentially all the resistance to int~rfacial

mass transfer of a sparingly soluble gas lies in the liquid film

"blo'v out" (spray formation) condition. The maximum liquid velocities that may be used tend to be quite low because of the long residence times typical of bioreactors.

The hydrodynamic regime of operation influences column performance. At low gas velocities (:s 0.05 ms- I

) in water­like fluids, bubbles h;l\"e spheroidal shapes and rise uniformly with no interaction. This is the bubble flo,,"' regime (Figure 20) which is also kno""n as "homogeneous" or "unhindered bubble flow". As the gas velocity increases the bubble motion becomes unstable and chaotic, and the column becomes more turbulent. Larger bubbles with little definition to their shape coexist with many small bubbles. This is the churn-turbulent regime (Figure 20). The transition from bubbly to churn­turbulent flow is gradual and it occurs oYer a range of gas flow depending on the properties of the fluids and on reactor geometry. Further increase in gas flow rates leads progress­ively to increased bubble coalescence, slugging and annular film flow; however, these other flow regimes are gener3lly

)t encountered in biore3ctor applications.

BUBBLE FLOW CHURN - TURBULENT

Figure 20. Bubble and churn flow regimes.

~as holdup.

t'he volume fraction of gas in dispersion, or "gas holdup" (E), is an important characteristic of bubble columns and other gas liquid reactors. By definition,

VGE = ---'=--- (59)

VG + VL

The reactor must be able to accommodate the gas holdup produced under various conditions. Furthermore, the resi­~n~e time of gas in liquid and hence the efficiency of util­IZatIOn of gas depend on gas holdup

hL E

. Fermenter technology

The specific gas-liquid interfacial are3 for mass transfer is controlled by a combination of gas holdup and me3n bubble di3meter

6E aL = ---- (61 )

dB (I - E)

where dB is the di3meter of a sphere h3ving the same surface­to-volume ratio as the gas bubble, In practice a distribu'tion of bubble sizes is encountered and dB is calculated using

N

I nj d/ dB = ,,-:i~~l~__ (62)

2L nj di i=l

",·here nj is the frequency of occurrence of bubbles with a diameter d i . Bubble size and frequency distribution may be measured, fOf example, by electrical resistivity and other similar probes. Alternatively, aL may be determined by one of the direct measurement techniques such as sulfite oxidation.

In bubble columns gas holdup shows the following general ) dependence on gas velocity (

E = a UGb (63) )

The parameters a and b have been found to e 2.47 and 0.97, and 0.49 an 0.46, respectIvely, for bubble (Uc < 0.05 ms I) and churn-flow (UG > 0.05 ms 1 regimes i, ai - for broad ranges 0 bub Ie column geometries. Addition of in­organic salts to water enhances gas holdup by a few percent up to an ionic strength corresponding to - 0.15 M NaCI. This effect is due to coalescence inhibition ",·hich results from electrical repulsion between like ions at the surfaces of bubbles. For any given ionic strength, the type of inorganic salt does not influence gas holdup.

Numerous other gas holdup correlations are aV3ibble in the voluminous literature on bubble columns (see Reading List). One example is the correlation7

;

(64)

where c is either 0.20 (non-electrolytes) or 0.25 (electrolytes). Equation (64) covers column diameters de = 0.152 - 0.6001 and U G = 0.004 - 0.33 ms-'. It applies to Newtonian media such as water, glycerol and methanol.

In homogeneous non-Newtonian systems the following equation may be employed

'-..-/

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mass transfer consumption term term

where qo, and X are the specific oxygen consumption rate (02/kg cells s) and cell concentration (kg celli m3)respectiyely. Interruption of the air supply to the reactor eliminates the mass transfer term (eq. 38) and the dissolyed oxygen con­centration declines linearly with time (Figure 14) due to

oxygen consumption by the biomass, The slope of the CL YS. t plot yields qo, X (Figure 14). The air supply is turned on before the dissolved oxygen concentration has dropped to the critical dissolyed oxygen leyel for the microbial species so that the fermentation is not damaged. The oyerall volumetric mass transfer coefficient is determined using the CL \'S. t plot beyond the point of resumption of oxygen supply. Thus eq. (38) is rearranged to

Air off

•

--'- _Ccrit L­

TIME Figure 14. The dissolved oxygen concentration (Cd vs. time.

CL = C,:· - _1_ (qO' X + d CL) (39)k L aL dt

and kLaL is obtained from the slope of a plot of C L YS.

(qO' X + d d~L) (Figure 15). Rapid-response dissoh'ed

oxygen probes should be used to minimize the effect of electrode delay on the measurements.

Gas-phase oxygen balance technique depends on measure­ments of mass flow of aeration gas into and out of the fermenter. The mass fraction of oxygen in the inlet and outlet gas ,streams must also be determined (mass spectrometer, paramagnetic oxygen analyzer) as well as the steady state dissolved oxygen concentration (dissolved oxygen electrode). The haL is obtained from the oxygen balance

M (xo - Xi) = V L h aL (C:' - Cd = V L qo, X (·W)

where VL is the broth yolume, M the mass flow of gas and x is the mass fraction of oxygen in gas (0 = outlet, i = inlet).

178

Figure 15. Calculation of kLaL'

Equation (40) assumes no eYaporation and it does not correct for carbon dioxide production; howeYer, the necessary cor­rections can be easily incorporated. With the steady-state method any possibility of affecting the fermentation by inter­ruption of air supply is circumvented.

i,;.

Heat Transfer

Most fermentations require careful temperature control. Heat generated by agitation and aeration power input and that generated by the fermentation itself needs to be estimated for design of sufficient cooling capacity. Sterilization operations also require knowledge of heat transfer and necessitate the provision of sufficient heating capability.

Typical fermentation heat generation for bacterial, fungal and yeast fermentations is of the order of 3-15 kW m -3.

The exact amount depends on the nature of the substrate and its rate of oxidation. A highly reduced substrate such as a hydrocarbon would release more heat per mole substrate on complete oxidation than a relatively less reduced carbohydrate:>-~) Methods for estimating the heat evolution have been discussed

in Bailey and Ollis.2

Between 1 and 15 kW m- 3 of heat input occurs due to

agitation in stirred tank fermenters. In bubble columns and airlifts the contribution of heat due to agitation is usually less than 5 kW m-3

• Once the heat transfer rate ("'hich equals the heat evolution rate pItts the heat generation due to agitation at steady state) is established, the heat transfer area needed to obtain this rate is calculated from:

(41)

where A H is the transfer area and /).T is the mean temperature difference dri,-ing force. U H , the overall heat transfer coef­ficient, is the sum of the resistances to heat transfer due to

the fluid films on either side of the heating or cooling surface, fouling (corrosion, protein burn on) resistances on either side, and the resistance due to the metal wall through which the heat must pass. Hence,

• •

••

velocity usin" most of the J\":tibbk correbtions is limited onl)" to specific reactor geometries 0\"([ narrow r:tnges of scale. Onl)" recently did a more general :tirlift design procedure become a\"aibble as discussed bter in this section.

The induced liquid circubtion is an important distinguishing characteristic of airlift reactors. In other typcs of bioreactars, such as the bubble columns and the stirred tanks, the general requirement of long residence times se\-crcly limits the maximum linear flow "c1ocit)" through the reactors unless recycle flow is en;ployed. In airlifts, high linear liquid velo­cities arc attainable without recycle and these lead ta improwd turbulence and good mixing, heat and mass transfer. The liquid circubtion in an airlift reactor originates from the difference in the bulk densities of the fl uid in the riser and the downcomer. The fluid circulates along a well defined path: upflo\\" in the riser, i0\\'nflow in the down comer. A mean circulation velocity (ULd is defined as

(74)

where Xc IS the circulation path length and tc is the average

t Air

External-loop oi rlift

. . . ,

o 0 •

Q 0 .. \)

, 0

• c

Internal-loop split-cylinder

air lift

Figure 22. External- and interaI-loop airlift reactors.

Fermenter technology

time for one complete circulation. The circulatory flo\'.' is clearly revealed by' injection of a tracer such as an acid pulse into the downcomer (or riser) and follo\\'ing the tracer flow at some downstream location. The characteristic deca:"ing sinusoidal tracer response depicted in Figure 23 is obsern:d: the time difference bct\\'een adjaceIH peaks is the circulation time.

Unlike an o\'eral1, awrage, circulation velocity (ULd, v~lues of a superficial velocity measured either in the dO\\'ncomer (ULd ) or the riser (UL,) are more me:lIlingful. The continuity criterion leads to the follo\\'ing relationship between the liquid velocities in the riser and the downcomer:

(75)

TRACER INLET

DETECTOR

Figure 23. Tracer response in an airlift reactor.

The superficial \'elocity must be distinguished from the "linear liquid velocity", also known as the "interstitial velocity", because in reality the liq uid flo\'.' occupies only a pan of the flow channel, th.e rest being taken up by the gas. The inter­stitial velocity (Vd and the superficial velocity are related as follows:

V Lr =~ (76)1 - f,

and

V Ld =~ (77)1 - Ed

The velocity of liquid circulation, \\·hile itself controlled by the gas holdups in the riser and the downcomer, in turn

l '---/

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.The mechanical power requirements of ungassed stirred tanks may be estimated using Power number (Po) vs. impeller Reynolds (Re;) number plots, examples of which are shown in Figure 17. The Power number and the impeller Reynolds number are defined, respectively, as

(-+6)

and

(-+7)

10 Flat-blade turbine

0 Turbulent Paddle0­

Prochem impeller 1.0

Marine impeller

0./ I 10 10 2 103 104 105

Re· (-)I

Figure 17. Power number vs. impeller Reynolds number for various impellers.

The exact nature of Po - Rei plots is dependent on impeller type and on the presence or absence of baffles. The power absorption by liquids in un baffled tanks in turbulent flow (Rei > 102

) is significantly less than in baffled tanks. In laminar flow (Rei ::; 10) the power number is inversely dependent on Rei:

Po :x Rej-t (-+8)

with the constant of proportionality dependent to some degree on the type of impeller. Under developed turbulent flow conditions (Rei > 104

) the power number becomes independent of the impeller Reynolds number, but depends on the impeller type. Because most applications are likely to

im'oh-e highly turbulent reactors, a compilation of constant PO""er numbers for yarious geometries is provided in Table 3. The values of the proportionality constant (eq. -+8) are :given in Table 4.

Introduction of gas into the mixing \"esse! always leads to a reduction in the power absorption relative to the ungasscd situation. Once an estimate of ungassed power (P) is available, the power input in the presence of gas may be calculated using the Michel-\1iller equation

180

Table 3 Turbulent Power Numbers in Stirred Vessels r

Geometry Po (Baffled tank) (-)

Propeller (square pitch, 3-blades) 0.32 Turbine (6-bladed) 6.30 Turbine (6-curved blades) 4.80 Flat paddle (2-blades) 1.70 Prochem impeller (5-blades, D i = dT /2) 1.0

P2 N D.3)O.45 P G = 0.72 ( Q0.56 I (49)

This equation provides a good approximation in many appli­cations but it should not be used for extreme values of gas volume flow (Q). Other design parameters such as the overall gas holdup (e) and the \'olumetric mass transfer coefficient (kLad can be calculated on the basis of available correlations ~.

which have been summarized by Mann. 6 Some useful correlations are 3

j 10 = 0.52 (N~i3f5 (PL ~ D f65 (~J 1.4 (50)

and

(51)

Numerous other equations are available which may be more suitable for specific situations.

Table 4 The Values of c in Po = c Rej-I for Various Impellers

c Impeller (-)

6-Bladed standard turbine (unbaffled -100 tank) Helical ribbon (unbaffled tank) -380 Propeller (3-bladed, square pitch, baffled -40 tank (4-baffles))

Non-Newtonian media

For non-Ne",·tonian media the impeller Reynolds number is based on the apparent viscosity of the fluid:

Rei = PL N D? (52) ~l.,p

For the often observed power law behayiour the apparent viscosity is giYen by

liquid flows in a YCrtica! pipe using either the' fluid of interest or a reasonJble' simuhtion of the fluid. The experimentation is yery simple and straightfonnrd. For example, for air­water the following applies

E =0 Uc, (85) r 0.24 + 1.35 (UCr + U Lry:'9J

when U Lr > 0.3 ms- 1•

Knowledge of the riser and the downcomer holdups enables the calculation of o\"(~rall holdup (c:):

Ar Er + Ad Ed E = (86)

Ar + Ad

and hence the height of gas-liq uid dispersion:

(87)

In equation (87) hL is the unaerated liquid height. Design at the hydrodynamic and mass transfer level would

ilvolye the prediction of U Ln £, En Ed and hD for any giYen operating conditions (thn fluids) and gi\'en reactor geometry (An Ad, Ab and hd. A design flow chart for internal-loop airlifts has been published. 12

Gas-liquid mass transfer

The measurement of the oycr:l1l yolumetric mass transfer coefficient and gas holdup in a gi\'en fluid in a small bubble column enables the calculation of the ratio kL/dB:

kL _ (1 - £) k a"'--_---'------'L=---=L = "4' (88)dB - 6E

This ratio has been found to be constant (ljJ) for any specific fluid oyer broad ranges of gas flow rates. Thus kL/ dB which has been experimentally determined in a bench-top model 'eactor may be used for estimation of kLaL in larger production vessels. An estimate of the gas holdup in the reactor is first obtained using the procedure described earlier; kLaL is calcu­lated as follows:

'\j.'6£haL (89)=0

(1 - £)

For air-water the parameter '\j! is ~0.053 S-I. For fluids made up of filamentous or fibrous solids suspended in a water-like medium "4!(S-I) depends on the concentration Cs (dry wt./vol.%) of solids:

(90)

where PL, flL, DL and a refer to the properties of the suspending fluid.

Other considerations

Substrate injection. The problem of location of substrate feed poin~s in an airlift vessel becomes p:lfticularly significant in COntmuous and fed batch operations. For rapidly utilized

Fermenter technology

substrates, the concentration of which must be kept low for reasons such as substrate toxicin' or substrate inhibition. the microorganisms in a tall airlift m'ay be stan-cd of the subs~rate only a shon distance dmvnstream of the point of substrate injection. Thus, multiple substrate feed points may be neces­sary axially up a reactor if product yield reduction due to substrate stan-ation is to be avoided.

The substrate balance for a differential volume of the riser may be written as

(91 )

where S is the substrate concentration at an,' vertical position z, VLr is the interstitial riser liquid w!ocity,'E Lr the riser axial dispersion coefficient of the liquid phase and .Rs the rate of substrate consumption. When the substrate concentration must not fall below a critical minimum value Smin and it should not exceed a maximum of Sma", because of inhibition considerations, then eq. (91) may be solved with appropriate reaction kinetic expression to determine the axial distance at which fresh substrate addition becomes necessary.

Gas sparger

Perforated plate gas spargers are often used in airlifts and, in keeping with the practice in bubble columns, these plates are located at the base of the riser in the airlift. However, this type of sparger positioning is inappropriate in airlift devices because the recirculating flow from the downcomer leads to a maldistributi'on of gas (Figure 24(a)). The use of perforated pipe ladder type gas spargers located just above the point where the flow from the downcomer meets the riser leads to

imprO\'ed gas/liquid flow (Figure 24(b)). Perforated pipes are recommended for bioreactor applications.

Immobilized Enzyme Reactors

Immobilized enzyme (and immobilized whole cell) catalysts (see Chapters 17 and 18) can be employed in a variety of reactor configurations. Catalyst panicles may be used in sus­pension as in stirred tank and fluidized bed reactors (Figure 25) or they may be held in place in fixed or packed bed devices. Hollow fibre reactors containing catalyst immobilized either throughout the thickness of the fibre wall or confined to one side of it (e.g. perfusion systems) are possible. Fbt polymer membranes containing immobilized catalyst ha\'e been used in spirally wound configurations. Immobilized particulate biocatalysts can, of course, also be used in airl;[t and bubble column reactors so long as the solids loading and density are not excessi\'e. In such reactors a compressed gas prO\'ides the necessary agitation in the fluid and gas-liquid mass transfer is not the main consideration.

Reactor efficiency is measured by the quantity of substrate tr:lI1sformed per unit time per unit mass of immobilized

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Table 6 A Comparison of Mechanically and Pneumatically

Agitated Bioreactors

Mechanical Agitation Pneumatic Agitation (Stirred Tank) (Airlifts, Bubble Columns)

1. Mechanically complex Mechanically simple and (stirrers, shaft, seals, robust bearings)

2. Often high shear Gentle, low shear levels (suitable for tissue culture, plant cells, fragile genetically

( a ) engineered microorganisms)

Prochem Maxflo hydrofoil 3. Gas throughput limited High gas throughputs by impeller flooding possible (particularly in

airlift devices) ---------------------- '---:­

4. Difficult to clean due to Easy to clean. Extended mechanical complexity; asceptic operation possible greater possibility of (useful in continuous contamination over operation) extended operation

5. Turbulence confined to More uniform distribution impeller zone in viscous of turbulence non-Newtonian media. Gas channels through the impeller zone while the rest of reactOr remains stagnant

6. Operationally flexible Limited operational (controlled by impeller flexibility. Require more speed and by gas flow careful design rate)

( b) Hydrofoil impeller

Sieve plates

fI • ~ ..

Vertical baff les-. . . - -: " . -.. ..­

-~-.

( c )

Air Air Air Figure 18. Some newer impellers. Figure 19. Bubble columns.

Intermig

--

Fermenter technology

.. • .. . 0

, D D 0 o " o • D' 0 Cot 01 yst " 0. t>

" • ~ " particles o D ..

• 0 ..

.. " 0.. " 0 ..0 .0

.. 00"

0

a) Stirred tank

b) Fluidized bed

..- L ­-- -<:::I":

W \..I; )1:

~

....,~ -y

+ Feed-.----"-----{ Feed

c) Packed bed d) Hollow fibre (e) Spiral wound membrane system module

Figure 25. DeploymerH of immobilized cat.11;'st: free suspension in stirred tank (a) or fluidized bed (b); fixed catalyst in packed bed (c), hollow fibre (d) or spira! wound membrane (e),

Fibre wall (permeable)

Membrane supporting cata Iyst

Product

form \vhich is indicati\'(: of equal performance of the two reactor systems in this regime. However, ""hen 5 « Ks the reaction rate is first order in substrate (eq. 92) and the plug flow system gives a b,>tter performance than the continuous stirred tank. In the lattcr, all the catalyst \\'ould bc exposed to a low substrate concentration and this can bc utilizcd ad\'anrageously in continuous stirred tanks ,,'hen the re:lction is inhibited by substrate.

The theor:tical efficiency of other types of reactors is between thc two extremes of the packed bed and continuous Stirred tank flow geomctries.

Mass Transfer Effects

Heterogeneous catalysis has its associated mass transf~r con­siderations. Mass transfer resistances at the interface of solid support and the bulk liquid and within the solid matrix often reduce the effectiveness of the immobilized form. Adnntages of immobilization should be weighed against possible disad­Yantages in the process of choosing a particular form of biocatalyst.

An:llysis of the interfacial and intrapar~c1e mass trJl1Sfer and cat:dyst performancc is illustrated for a spherical catJ1:.st

_~:k1--;-=,,- GAS

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The shear rate (1') expression commonly employed for the calculation of apparent yiscosity of fluids in bubble columns (sec eq. (3» is

l' = 5000 UG (66)

which is due to Nishikawa and coworkers.9 This expression [eq. (66)] is used for the calculation of ~Iap in eq. (65). Howeyer, there is a considerable degree of uncertainty on the mean shear rate in bubble columns.

According to some recent work, to the simple holdup equation (eq. 63) should apply to non-Newtonian media also. The parameters a and b now depend on the properties of the fluid as well as on the flow regime. The parameter b has been empirically correlated with the flow index according to

b = 0.564 n -0.354 (67)

Equation (67) disregards any flow regime effects, but It IS

based on data on a variety of fluids including fermentation broths of fungi Chaetomium cellulolyticum and Neurospora sitophila. Other gas holdup data on slurries which simulate fungal media is available elsewhere. ll ,12

Gas-liquid mass transfer

T\\·o of the correlations for the overall yolumetric mass transfcr coefficicnt in N cwtonian fluids arc:

and

(69)

These equations were developed by Fair13 and Akita and Yoshida7

, respectively. For air-water, a simple cquation is

kLaL = 2.39 X 10-4 (PG lVd·S6 (70)

which has been shown to apply up to a height-to-diameter ratio - 24. Notice (cqs. (68)-(70» that the oyerall yolumctric mass transfer coefficient may be based either on the liquid yolumc (kLad or on the volume of gas-liquid dispersion (kLaD). These two arc related as follows:

kLao = kLaL (1 - E) (71)

The: mass transfer ,york on non-Newtonian media in bubble columns is less extensive. Some equations which may be useful in estimation of mass transfer performance are

kLao = 8.35 X 10-4 U 0044G

II -1.Q1•.tp (72)

due to Godbole et al. s and

kLao = 3.15 X 10-3 U G 0.59 '[ -0.S4,.lp (73)

J due to Deckwer et al. 14

• For additional information on non­Newtonian systems the work of Schumpe and Deckwerl5

rshould be consulted. Gas-slurry systems haw been treated 11.16

elsewhere. A Yast amount of literature on bubble columns is available;

some of the main sources are listed in the Reading List.

Airlift Bioreactors Airlift bioreactors consist of a liquid pool divided into t\\·o distinct zones only one of which is usually sparged by gas. The different gas holdup in the gassed and ungassed zones results in different bulk densities of the fluid in these regions which causes circulation of fluid in the reactor by a gas-lift action. The part of the reactor containing the gas-liquid upflow is the "riser" and the region containing the downflowing fluid is known as the "downcomer". Figure 21 shows the schematic of an airlift reactor.

.

'-..t->=t-- DOWN COMER (DOWNFLOW)

SPARGED RISERS (UP FLOW j

~---GAS SPARGER GAS ....-+~~+-'+--:-'. ~

Figure 21. Schematic of an airlift reactor.

Airlift reactors have been successfully applied to almost evcry type of fermentation. Many examples haye been cited

17in other works.5• Recent applications include hybridoma

cell culture for monoclonal antibody production on a commercial scale.

Airlift reactors are ayailable in two basic forms: (i) the internal-loop airlifts in which the riser and the downcomer are contained in the same reactor shell, and (ii) the external­or outer-loop reactors where the riser and the downcomer are two quite separate tubes which are linked ncar the top and the bottom. The external- and internal-loop configurations are shown in Figure 22. Modifications to the basic airlift dcsign ha\'e been used to produce othcr sub-types of airlift reactors, some of which have been discusscd by Chisti and Moo-Young. 5

Estimation of such essential airlift reactor design parameters as the overall gas holdup (E), volumetric mass transfer coef­ficient (kLad and the magnitude of induced liquid circulation

Bioreactor Scale-up

Laboratory scale bioprocess denlopment identifies the opti ­mal fermentation conditions for the process. Oxygen transfer requirements, maximum tolerable le\'els of shear, pH and temperature control needs should become known at this point. The object of scale-up is to reproduce on pilot or production scale the successful fermentation results achiend in the lab­oratory. The results are often specified as prod uction rate per unit fermenter volume.

In practice scale-up is quite complex. It is not generally possible to reproduce exactly on the production scale all the various parameters from laboratory or pilot scale units. For example, at equal specific power inputs twO geometrically similar stirred reactors do not have identical mixing times. As a result scale-up is based on the strategy of holding constant only one or two of the several possible parameters at different fermenter scales. The parameter(s) held constant are those which are considered to han the greatest impact on the fermentation; furthermore, the criterion of geometric simi­larity (i.e. keeping the ratios of corresponding lengths equal on production and pilot-scale units) is not always rigidly adhered to so that small geometric variations may be utilized to advantage as long as they do mot result in unpredictable behaviour.

The scale-up methods which han been most often proposed are as follows:

1. scale-up based on equal power input; 2. scale-up based on equal mixing times; 3. scale-up based on equal oxygen transfer (kLad; 4. scale-up based on equal shear rates (or impeller tlp

speed).

The list is not exhaustive. For highly aerobic fermentations ~cale-up based on maintaining a constant oxygen transfer rate IS a reasonable approach but in other fermentations, limitations such as those on shear rate may be equally important.

The following comments on scale-up apply to stirred tank trpe of fermenters. Considerations for scale-up of pneumatic reactors, particularly the airlifts, "-'ere examined earlier in this chapter.

Scale-up based on equal PCIVL ratio

Th.e criterion of equal PG/V L on pilot-plant and production unltS has been employed for certain antibiotic fermentations. The. available evidence indicates that the necessary power reqUIrements decrease with increasino fermenter volume ap­

. b

prOXimately as

PG <X V -0.37 L (112) VL

Consequently, keeping PG/V L constant in scale-up may not

Fermenter technology

be :In energy efficient approach. Furthermore, it may not be ' a satisfactory strategy for shear sensitiYC fermentations since the impeller tip speed and Rernolds number scale-up by factors of > 1 when PdVL is held constant for geometrically similar nsscls. Table 8 shows some of the effects of geometri ­call:' similar scale-up of a 20 L reactor to 2.5 r:13 plant \"esse!. Effects of keeping PG/VL constant (i.e. PG/VL == 1 arbitrary unit for both reactors) on impeller rpm (N), tip speed (:\0,) and Reynolds number (Rej) is shown in Table 8. These par­ameters scale-up by the respective ratios of 0.34, 1.7 and 8.S. Table 8 also shows the~feets of maintaining constant rpm (N), constant tip speed (N OJ) and constant impeller Reynolds number (Re;).

Table 8 Effects of Scale-up Based on Constant PdVL (or Constant

N, N OJ, Rej) on Other Parameters

Laboratory reactor Plant reactor

Parameter 20 L ·2.5 mJIi / V,

~G/VL 1 25 0.2 0.CJ16 N 0.34 1 0.2 0.C4 N 0i 1.7 5 1 0.2 Rei 8.5 25 5 1

Scale-up based on equal shear

The maximum shear rate is related to the impeller tip speed which is held constant on scale-up. However, the shear rates in the fluid which are governed by fluid· turbulence or Reynolds numbers do not scale-up proportionately because the impeller Reynolds number does not remain constant. (Table 8).

Scale-up based on equal oxygen transfer

Maintenance of an equal o\-erall volumetric oxygen transfer coefficient (kLad is often taken to ensure equal oxygen transfer on scale-up. Ihis is t~ue only when the oxygen transfer dri\-ing force alSo remains unchanged on scale-up. - For stirred tanks

(113)

where k] is dependent on geometry and k2 and k J are scale­dependent.

More complex scale-up methods rely on estimation of kLaL as well as the spatial oxygen concentration profiles in the reactors to yield a value of oxygen transfer rate. Operating and scale-up parameters are adjusted until the desired transfer rates are obtained for realistic operating conditions.

(

I

·\f

Biotechnology/The Science and the Business

affects these holdups by either enhancing or reducing the \'elocity of bubble rise.

Airlift reactor design

Probably the first question faced by the designer of airlift bioreactors for a particular application would be one of choice of configuration: external-loop or the internal-loop. Table 7 compares the performances of these two distinct geometric types of airlifts; such a comparison could form the basis of a preliminary choice. Generally internal-loop reactors have better mass transfer characteristics. On the other hand, in fluids in which \'Cry high viscosities necessitate greater tur­bulence and ~hear for adequate mixing and. mass transfer, the external-loops may be preferable.

An energy balance over the circulating airlift loop can be used to obtain the following equation for the superficial liquid \·elocity in the riser: 18

(78) __T::---= + K r

(1 - E r )2 B Ad (1

= [ K 2g h (E(rA-)~d)U Lr D

Equation (78) is for low viscosity water-like fluids and it ignores wall friction losses in the riser and the downcomer. The equation applies to external- and internal-loop configur­ations of airlift reactors.

The parameters KT and KB are the frictional loss coefficients for the headspace and the bottom of the airlift reactors. For typical internal-loop airlift the term containing KT (eq. 78) can be ignored, while KB is dependent on the bottom geometry:18

(AAd)O.8KB =ll.4 (79)

b Equation (79) applies over an Ad/ Ab range of 0.2 -1.8; Ab is' the free area for flow between the riser and the downcomer. In external-loop reactors K B = KT and a KB of 5 may be

Table 7 Relative Performance of External- and Internal-Loop Airlift Bioreactors

Reactor

Parameter External-loops Internal-loops

Mass transfer (kLad Overall holdup (E) Riser holdup (lOr) Downcomer holdup (fd) Liquid velocity (ULr) Circulation time (tJ Liquid Reynolds Nos. (shear) Heat transfer

r86

lower lower lower lower higher lower higher probably higher

used for design purposes for the following approximate geo­metric ranges: Ab/Ad = 1-2, Ab/Ar = 0.25-1 and Lcp/d,p = 2-7.

Recent research19 has shown that for non-Newtonian, pseudoplastic fluids, for which eq. (78) is unsuitable, the following may be used for U Lr calculation:

ULrAr (~PFr + ~PFd) - PL g hD ULrAr (lOr - Ed)

1 3 [KT (A r )2 1 ]+ 2 PL U Lr Ar (1 _ E )2 + KB Ad (1 _ Ed)2 = 0 r (80)

The ~PFr and ~PFd in this equation are the frictional pressure drops in the riser and the downcomer, respectively. The in­depth procedure for the determination of these pressure drops is described elsewhere. 19 Equations (78) and (80) assume that the riser and the downcomer gas holdups are known; these parameters are interrelated:

(81)

for internal-loops without a gas-liquid separator per se, and"-.-/

Ed = 0.79 lOr - 0.057 (82)

.Ed = 0.46 lOr - 0.024 (83)

for external-loop airlifts. Equation (82) iSjSuitable for water­like fluids while eq. (83) is more apprdpriate for slurries encountered in fermentations of such fungi as Penicillia and Aspergilli and in the cultivation of filamento~smicroorganisms like Streptomyces. .

The calculation of riser holdup for re~ctor design and scale-up requires some experimental investig~tion,particularly when new applications are involved and th~ fluids arc rheol­ogically complex. Equations of the type ­

need to be established by independent ,!~riation of gas and

higher higher higher

~~~he:r "\~'.

higher lower

probablY lGlwer "

r

J.

I~

i

I

I ~ I

OverflOW

IJ Solids

outlet

~_Slurry

Tubular bowl

Fermenter technology

aeometlT . The v:tlue of L represents the are:t of a graYity ;ettling ~:tnk which is capable of the same sep:tr:tting ability for continuous flmY oper:ttion as the centrifuge. For ap­

ropri:tte equipment selection the separation requirements h:tye to be defined. It is uneconomical to specify equipment for more stringent st?p:tration duty th:lI1 is rt?ally nect?ss:try. Cle:trh', from eq. (1 H) the particle diameter and the density differ;nce between the particle and the suspending fluid are

important factors whi~h affect separation. . . Selection of a centrIfuge for any ne"'" appbcatlon ""'ould

almost always in\'oh"e expensi\"C pilot scale e\·aluations. A few simple laboratory tests can, hO""'e\'er, prO\'ide an indication of whether or not the pilot run is e\"Cn worth pursuing, If a sample of the slurry does not settle on standing under the influence of gra\'ity o\"Cr seyeral days, it is unlikely that a separator can achie\"C yery ~uch., Howe.ve{', change in s~mple characteristics such as particle sIze ""'hICh may be achle\'ed, 'for example, by the addition of flocculating agents or by alteration of pH may implO\'e the likelihood of separating difficult to settle solids, The flocs should be strong enough to withstand the accelerational forces which are experienced as the fluid enters the centrifuge and comes up to the same rotation:tl speed as the bulk of fluid. If the flocs formed are easily disintegrated there is little ad\"Jntage to adding floccu­lation chemicals.

A slight shaking of a bottle of gravity-settled solids can provide an indication of how light the particles are. Other solids properties such as the particle size and density need also to be known. Here the density refers to solids as they are in suspension and not dry. This difference is of particular importance for biological solids which contain a high pro­portion of ",,'ater and s""'ell when more is added. The solid s may be fibrous (fungal mycelia) or slimy, or may occur as pellets. These properties impact upon the choice of centrifuge. ~or example, some fibrous materials settle as mats under the high centrifugal force and may cause de-sludging problems in certain automatic solid-discharging centrifuges. Solid packing characteristics and ease of settling can be easily judged by spinning a test tube sample in a laboratory centrifuge at - 3000 rpm for 3 to 5 minutes,

Equipment

SeW!"J! types of centrifuges arc a\'ailable; the more common ones are:

(i) tubular bowl; (ii) multichamber bO""'I;

(iii) disc-stack; (iv) scroll discharge decanter centrifuge.

The particle size ranaes for which these configurations are• t> ~

suItable are shown in Figure 28.

Tub/dar bowl centrifllge. Shown schematically in Figure 29, the tubular bowl is the simplest centrifuge configuration.

I ~ I ~

~ I

~i I

~

basket .ba I I!I Scroll discharge ,

I

I

Ultra II I 1II

I Tubular bawl

I Batch disc I

NOlzle disc

I ~ Valve disc

§ ~-=@ Opening bawl

I d ImperforateI I

I I

I I I

10-2 10- 1 /0 102 103

PARTICLE DIAMETER (fLm)

Figure 28. Panicle size range for different types of centrifuges.

--., ,-Liquid

- C ~RetalJ1ed~ 1 solids

SOlids

~r'~ ""'"u."~Tt:,} (~.. , Retained

solids

Decanler bowl

Figure 29. Centrifuges.

High 'g' -forces do permit good solids dewatering but the operation is batch with respect to solids. The solids-handling capacity is limited and solids recovery is labour intensive. Consequentl~" only slurries with low solids concentrations can be economically de""':ttered.

flultichmnber bo'wl centrzfuge. This configuration (Figure 29) is basicall:-' a tubular bowl centrifuge with increased solids handl~lg capacity. Efficiency is maintained up to complete filling of the chambers. Other oper:ttional char:tetcristics arc nearly the same as for the tubular bo""'] machines.

Disc-st"ck centrzfuges. Disc-stack centrifuges come in sewral types depending on whether the solids are retained or dis-

Mullichamber bowl

Disk slack

I i Biotechnology/The Science and the Business

DOWNCOMER

t GAS

(a )

t GAS

( b) Figure 24. Positioning of gas spargers in airlift reactors: (a) poor gas distribution; (b) irnpro\'ed gas injection.

catalyst for specified initial concentration of substrate (50) and its desired conversion X = (50 - 5)/50, The conversion characteristics of different reactor configurations can be calcu­lated from a knowledge of the kinetics of reaction. Thus, for a reaction which obeys lvlichaelis-Menten kinetics

d5 kr e 5 (92)

dt - Ks + S

(e enzyme concentration, Ks = Michaelis constant, 50 substrate concentration), we have for various reactors:

Bateb stirred tank

Change in quantity of Rate of substrate

substrate in the reactor consumption by reaction (93) or

. d5 kr E 5 (94)- VL dt = K + 5s

where E is the total amount of catalyst in the reactor. Equation (94) may be integrated for 5 = 50 at t = 0 and 5 = 5 at t = t,> to

kr E t = 5 _ 5 _ K In i (95)sV L 0 50

which can be rewritten in terms of conversion X as

k E t - ­~ = X 50 - Ks In (1 - X) (96)

Continuous stirred tank ~ ::­

The substrate concentration in the inlet stream is 50 and because the reactor is well-mixed the substrate concentration .. in the exit stream (5) is the same as in the volume of the '---/ reactor. A steady state substrate balance in the reactor can be written as

5ubstrate flow into reactor = substrate flow out of reactor + substrate consumption due

to reaction (97) or

Q 5 o

= Q 5 + Kr E 5 Ks + 5 (98)

which can be rearranged and written in terms of X:

krE Q

= 5 0

X+ KsX 1 ­ X (99)

Packed bed

Following the procedures outlined in the earlier examples, the appropriate equation for a packed bed system with feed flow rate Q is

krE - IQ = 50 X - Ks n (1 - X) (100)

Because Q is the volume processed in time t in a continuous flow reactor and V L is the corresponding volume for a batch reactor, comparison of eqs. (96) and (100) shows that the performance of batch stirred tank and plug flow systems is identical. This is a general conclusion, irrespectin of reaction kinetics. Howe\'er, kinetics alone do not determine reactor choice and operational considerations are important. For example, control of pH is operationally easier in stirred reactors.

for reactions which display Michac1is-i\1enten kinetics, continuous stirred tank and packed bed reactors operated such that 5 » Ks, eqs. (99) and (100) reduce to an identical

188

JIil"'"''III'--------- ­

1 dV L ,0.P (11 S)

Ar~ = hRF

The flow resistance Rr is the sum of the resistallce due to

nlter medium (rn,) and that due to accul11ubted biomass:

. \\' Rr = 0: - + r (119)A mr

where c\: is the mean specific resistance of the biomass cake and w is the dry weight of accumubted biomass. At constant pressure (,0.P) plots of tlV L vs. VL for incompressible filter cakes are linear, with slope and intercept dependent on con­ditions of operation. However, biological materials usually produce compressible cakes and experimental determination of filtration volume vs. time relationship is necessary.

Microfiltration and ultrafiltration

Microfiltration and ultra-filtration rely on porous membrane filter media. The basic difference between the two operations is the "particle" size range handled, Microfiltration mem­branes retain suspended solids down to ~O.05 ~lm. BJCteria, y~asts, fungi and tissue cells are readily removed while pro­tell1s and enzymes pass through the filter membrane at high flux. Ult,rafiltration membranes haw much finer pores (1- 20 nm) which allow retention of proteins, enzymes and car­bohydrates of nrious molecular weioht cutoffs. The followin o discussioll places emphasis on the ~icrofiltration of bacterial cells. Note that industrial ultranltration and microfiltration srstems are physicallJ' and operationally simibr; the theore­tICal funda~le~tals of these two operations are equinlent. A recent re\'le\\~o on ultrafiltration should be consulted for ~urther details'. In practical processing operations microfiltration IS usually employed in a cross-flow mode. The fluid to be filtered flows parallel to the filter surface (Figure 31). The :ross-flow of f:ed with respect to the filtrate flux generates shear forces which help to sweep the filter surface of excessive solids build.up. However, in most cases the buildup of a thin layer of sollds (concentration polarization) cannot be entirelv prevented. '

The filtrate flux (]) through the membrane depends on the ~rans.me~brane pressure ,0. Pn,i> the viscosity of the suspend­mg llquld (~ld and the hydraulic resistance of the membrane

Filtrate flux

Feed _ • • . • . I .... •:'.:: :.- +.<. ,,0, • _!(....,o~.:;.;.:o:o~."O72"3:=<;<3:<3<:: I

Retenfate (concentrated slurry of particles 1

Filter membrane

~ Membrane support

Filtrate or permeate

Figure 3 I. Principle of cross-flow filrer.

Fermenter technology

(R:-r) and the deposited solids (Rc):

] = ,0.Pn1 (120)

~lL (Rc + R:- 1)

The resistance of the solids layer depends on its thickness (65), voidage (£5) and the particle size:

180 (1 - (5)2 65R = -~-_.::.!..---.::.

dp £53c 2 (121 )

Biological solids from compressible cakes: the porosity (£5) of the solids layer decreases with increasing transmembrane pressure and the flux does not increase linearly with ,0.Pn1 , A typical rebtionship between filtrate flux and the trans­membrane pressure is depicted in Figure 32, While for pure liquids the flux varies linearly with ,0.Pn l> for slurries in­creasing ,0.Pn1 beyond a certain point produces no additional benefi t. As shown in Figure 32, for fixed ,0.PTM and other operating parameters, an increase in cross-flow velocit\· en­hances the permeate fl ux due to red uced solids byer thickness at higher flow rate. Reliable theoretical prediction of the behavi~ur ~f a solid deposit is not possible and experimental enluatlon IS necessary to determine suitable operatino con­ditions. The flux usually increases with temperature (r:duced viscosity) and with increasing flow rate parallel to the mem­brane. In biological applications the upper temperature limit would be determined by consideration of product stability (usually < 40°C). The choice of cross-flow rate would be a balance between the pumping costs and the higher filtration rate. The flux depends also on the concentration of solids in the bulk fluid since it affects not only the rate of transport of deposited solids back into the bulk flow (hence os) but also the turbulence intensity on the retentate side. The flux de­creases with solids concentration in the feed.

The process equipment consists of polymer membrane filters mounted in various ways. Ceramic membranes are no\\' a\'ail­

/ /

/ I.

I/ Pure liquid flux

/ /

/ ~_----U2/

/ t increasing / cross-flow

/ _------ u, veloci ty

TRANSMEMBRANE PRESSURE (6PTM )

Figure 32. Typical relarionship between filtrate flux and trans­membrane pressure.

195

Biotechnology/The Science and the Business

particle. For a substrate (e.g. oxygen or glucose) diffusing into the catalyst particle the general substrate balance on the particle is

Rate of diffusion into particle = rate of consumption (101)

since at steady state there is no accumulation. Solution of the appropriate balance equations leads to

(102)

where 5 is the substrate concentration at any radius r in the spherical particle of effectiYe diffusivity Dc; Rv is the \'olu­metric reaction rate. Equation (102) may be solved, either analytically or numerically, using the applicable kinetic ex­pression (RJ to obtain the positional yariation of substrate concentration inside the particle. For example, for a first order reaction with rate constant kn the expression for con­centration at any radius r in the catalyst bead is

5 Rp sinh [(k/Dc)O.5 r] S; = -; sinh [(k/Dc)o.s R ] (103)

p

Equation (103) applies in the absence of interfacial mass transfer limitation, i.e. when concentration in the liquid at the solid/liquid interface (Sj) is the same as the bulk liquid concentration. The total rate of reaction in the particle (RvVp ) is obtained by equating the total rate to the total diffusiYe flux at the surface of the particle:

_ 2 D dS I (104)RvVp - -4rr R p c dr R p

dS/dr at the surface being calculated by differentiation of eq. (103) followed by replacement of r with Rp • The diffusional resistance in the particle giyes rise to a concentration profile within the particle, the intraparticle concentration being less than in the bulk fluid. Hence, the average volumetric reaction rate is lower in the catalyst compared with a homogeneous bulk liquid reaction. The ratio of the reaction rate obserwd in the catalyst (presence of diffusional resistance) to the hy­pothetical rate for the same reaction in the absence of diffusional mass transfer resistance (i.e. liquid phase reaction) is the catalyst effectiveness factor YJ:

Rate of reaction with mass transfer limitation YJ = (iOj)

Reaction rate in the absence of mass transfer limitation

For the example of a first order reaction, the maximum possible reaction rate (no mass transfer limitation) is

(lC6)

Generalized plots of the effectiYeness factor (11) for any nth order reaction can be found in the literature. Figure 26 shows the effectiYeness factor plotted against a Thiele modulus <;>

19°

whi~h for any particle shape (sphere, cylinder, slab) is

I<P = Vp (~k St- )112 (107)

A 2 Dcp

when n > -1. Figure 26 is for a spherical particle and first order reaction (n = 1). Clearly, for spherical geometry and first order reaction the mass transfer limitation is negligible (i.e., 11 = 1) for 0.3 :s <P (Figure 26).

The effect of mass transfer external to the particle (solid­liquid interfacial mass transfer) on its effectiwness needs also to be evaluated. Use of the boundary conditions

dS ks at r = R - = - - (5 - 5) (lOS)

p' dr Dc I I

dS at r = 0-= 0 (109)

, dr II,

-­in the mass balance equation (eq. 106) leads to an equation I which includes the external mass transfer. The mass transfer l coefficient ks is calculated from the well-known correlations, applicable to a particular hydrodynamic regime. For example;-----~ for stagnant fluid around a spherical particle (i.e. negligible I difference between the density of the particle and that of the I suspending fluid), we haye

1.0 i 0.8 0.6

7](-) 0.4

0.2

, 0.2 2 4 6 8 10

_____Ji

Figure 26. Effectiveness factor (11) ys. Thiele modulus (<I» for II

spherical particles and first order reaction.

ksd pSh = D = 2.0 (110)

L

where Sh is the Sherwood number and for forced eonyection, the Fr6ssling equation

k·d .33Sh = ~ p = 2.0 + 0.552 Reo.s SeO. (111)

L.

where Re and Sc at;i-'.-the Reynolds and the Schmidt numbers, respectively. .'/

Seyeral other' equations for k s are a\'ailable for particular applications such as particles in fluidized beds and slurries in stirred tanks.

)

Retention sieve

......--Coolont

'---

r-++--- Retention sieve

C:::=:=J-./H----- Agitator discs

t Feed

Figure 34. Bcad mill (vertical chamber).

rRotaling disc

Aglta.ors ,Feed

/;Bearing seals

... I

I / ~

,"--------.- -----.J I I L J Coolant

Figure 35. Bead mill (horizontal chamber).

machines tends to fluidize the grinding beads to some degree thereby reducing grinding efhciency.

TIle kinetics of cell disruption in bead mills depend on the construction of the mill. First-order disruption kinetics haw been found in mJchines with predominJnt plug flow, whereas in machines in which the rotor design (Figure 36) permits signihcant backmixing the disruption de\·iJtes from the first order behJviour. For hrst order bJtch disruption the rate of protein release by cell rupture is directly proportional to the amOUnt of unrelc:ased protein:

dF.. . . dt = kD (R,n - R) (122)

where R is the weight of protein released per unit weight of cells, and Rm is the maximum measurJble release of protein. Equation (122) may be integrated for batch time (t) to give

Fermenter technology

Figure 36. Some rotors for bead mills.

In (. F..m .) = In D j = kDt (123)

Rm - R

where Di is the reciprocal of the fraction of unreleased protein. For continuous disruption in mills in which the flow may

be described in terms of the continuous stirred tank in series (CSTR) model the disruption kinetics follow the equation

- Rm- (k ;·)iJ (124)D i - . . - [1 + Dt r )

Rm - R

where tr is the mean residence time and j is the number of CSTRs in series. The \"alue of j may be obtained experimentally from residence time distribution studies, whereas tr is given by

V c t r = Q (125)

The disruption rate constant, ko , is a function of seYeral parameters: temperature, impeller rotational speed, bead load­ina

0' bead size and cell concentration. In addition, the densit),

of bead material is expected to affect kD although little has been written on this subject.

\Vithin limits, the disruption rate constant, kD , Increases with the agitJtOr tip speed, U T :

kD = KyU, (126)

the practical upper limit on the impeller tip speed being -15 - 16 ms -1. The po"·er consumption increases with agita­

tor speed according to

p ~: C NY diS (127)

where Nand d i are the rotation speed (rps) and the agitatOr dise diameter, respectively. The constant c is a function of the agitator design, suspension viscosity and the density of suspension. Heat production and the associated cooling re­quirements increJse with increasing agitJtion as also does the weJr on the beads, agitator and the chJmber WJlls. A certJin amount of weJr is unavoidable and bead material must be cJrefully selected in cell disruption applications. Beads Jre JYJibble in vJrious types of gbss, cerJmics and steels; note thJt some materiJ]s (leJded gbss, for example) will not be acceptJble with products intended for phJrmJceuticJI or food applications. The kinetic energy of the beads is an important £Jctor in disruption; it mJy be calculated as

Biotechnology/The Scienceand the Business

DO\VNSTREAM PROCESSING OPERATIONS

General Considerations

The output of a bioreactor usually undergoes one or more downstream separation, purification, stabilization and packag­ing operations to produce a saleable product. This section examines some of the more common unit operations employed in downstream bioprocessing.

a.-erall process design inyolYes consideration of interactions between various downstream and upstream stages and the bioreactor. In general, the smaller the number of processing steps necessary the more attractiye the process.

Solid-Liquid Separations

All bioprocesses involve one or more solid-liquid separation steps. Following fermentation the biomass may need to be separated from the broth, or the cell debris may have to be remoyed following cell disruption. Purification of enzymes and biochemicals by precipitation and crystallization also employs solid-liquid separation processes.

The well established ~olid-liquid separation technologies such as sedimentation and filter presses are not treated in this section even though they find numerous applications in the biotechnology ind uStry. Excellent treatments of these subjects

f can be found in the chemical engineering literature. Here \ only the operations of particular, interest in bioprocessing are \ examined.

Centrifugation

Adyances in structural steels have made possible the use of high-speed, corrosion resistant centrifuges for large scale bio­processes. Attention to ease of cleaning, containment and aerosol suppression, and the availability of sterilizable ma­chines have made centrifugation a yery important separation operation in biotechnology industry.

Operational principle

A centrifuge is basically a sedimentation tank with enhanced graYitational force to increase the rate of sedimentation. A particle enters the cylindrical bowl of the centrifuge (Figure 27) ,~·ith the feed flowing at some constant yclocity. The centrifugal force driyes the particle outwards towards the walls of the bowl. A particle initially at radius ri (Figure 27) would be at some position r after time t. If the residence time of the liquid in the centrifuge is such that the r 2: rB, the particle will reach the wall of the centrifuge and it will sediment. For particles in Stokes' law regime (Re «1), the terminal settling Yclocity at r3dius r is P

!?xiS of rotation

Zs

I

I LIQUID

~r~t ~ I

\,\ Path of'J

, I particle \ \ \ \

~ Lrs-!

t Feed

Bowl wall

Figure 27. Particle motion in the bowl of a centrifuge.

{J}r (Ps - pd d/U l = ( 114)

18 ~lL

where w is the angular YClocity of the centrifuge. Since U l == dr/dt, equation (114) may be rewritten and integrated between the limits r = ri at t = 0 and r = rB at t = t r (the residence

!time) to give

l (115) I

~L Note that ri cannot be zero and the feed must enter the"'--- 1

separating zone some distance from the axis of rotation. The acceptable yolume flow rate through the centrifuge

which will allow the particle of diameter 2: dp to separate may be calculated as

\T' d'?? = -.!: = w-(Ps - pd p-;r (rB- - rj-) ZBQ

t r 18 ~lL In (rB/r j ) (116)

The design of the centrifuge and strength of construction materials limit the maximum rotational speed.

Performances of different centrifuges are compared on the basis of the sigma concept which is also the commonly used basis of scale-up. The parameter L is defined as

L=~ (117)2 U l

Because U l (eq. 114) is dependent on the geometry of the centrifuge (yi3 ZB, rn, rj), L (m2

) is dependent on centrifuge

192

piZ0'Z

disruption rate constant is sensitin' to the design of the valve seat and some designs are shown in Figure 38. While most of the available disruption data applies to nh'e scats type (a) in Fi"ure 38, recent work has shown that vain.' scat (c) is

'" significantly bener than other designs in cell disruption

applic:nions. From equation (131) it can be seen that the operating

pressure is the major influence on disruption rate. Typically the operating pressure does not exceed 50- 60 MPa but higher pressure (- 130 1\lPa) equipment is anilable. An optimal choice of operating pressure is important because the power consumption during disruption is a linear function of the operating pressure, corresponding to about 3.5 kW per 100 MPa of operating pressure. The operating pressure also affects heat generation: at operating temperatures higher than 40°C, protein denaturation during disruption may occur. Since the temperature rise across a homogenizer due to adiabatic com­pression is about 2"e per 10 MPa, inadequate precooling, or failure to cool between multiple passes, can result in tempera­tures above 40°C and consequent denaturation. Although the

degree of disruption is increased by the number of passes through the homogenizer, a minimum number is desirable in practice. Multiple passes not only reduce the machine throughput but also cause further disintegration of already broken cell debris \\'hich leads to separation problems further downstream.

A "microfluidizer" high pressure homogenizer which relics on complex interactions between multiple liquid jets, cavitation and impingement has recently become a\·ailable. For bacteria such as E. coli and B"o'Uus Sltbtilis this device gives a perform­ance similar to the more traditional homogenizers but 95% breabge of yeast (S. cerevisiae) required 30 passes in one case compared with a residence time of only 3.3 minutes in a bead mill.

Other considerations

A unit operation cannot be considered in isolation from the rest of the process and the overall process must always be kept in mind. For example, the cell disruption operation afkcts the physical properties of cell slurry such as viscosity, density, panicle size and settlability of suspension which in turn affect the subsequent processing.

(a) ( bl (c) (d)

Flal type Knife edged Cane Iype Groo>ed

Figure 38. Valve seats for high pressure homogenizer.

Fermenter technology

Precipitation

Purification of enzymes and other biologically active proteins is often a multistep process. A single purification may in\'olyc cell disruption, debris removal, fractional precipitation, ion­exchange chromatography, gel filtration, affinity chromato­graphy and crystallization. The smallest possible number of separation steps is desirable for reasons of economy. Generally, no more than seven are used; many industrial processes make use of relatively crude preparations of enzymes and one or t\\·o purification stages may be sufficient. The purification scheme may be configured to produce several products from a gi\'en mixture so reducing the unit COSt of purification.

Selective precipitation of proteins from a solution of se\'eral proteins is among the oldest of purification and concentration techniques. Purification factors of 3-10-fold are reiatiYCly modest compared with chromatographic methods_ Howeycr, precipitation methods can deal with large quantities of material which may be quite crude (cell debris, suspended solids, contaminants). Furthermore, continuous operation can be cffecti\'e!y employed. Precipitation is encountered frequently in protein purification schemes, predominantly as one of the early purification stages. Additional steps may be used' downstream to polish further the product obtained at the preCipitation stage.

The process of precipitation converts the soluble protein to an insoluble form by altering the solute-solvent interactions. Protein molecules carry positive and negative charges, the net charge being dependent on the solution pH, At the isoe!ectric pH the protein molecule carries zero net charge and is least soluble in a pobr soh'ent such as water. The isoelectric pH of different proteins is usually different. Hence, by stepwise variation of pH of a protein solution different protein fractions may be precipitated and collected but exposure to extreme pH values may denature proteins and cause the loss of their biological activities. Protein precipitation by pH variatiOn is employed in the food industry (e.g. in milk coagulation).

Unwanted, heat-labile protein may be coagulated out of solution of relatively thermally satable components by heating.

\'\'ater is a strongly polar solvent (with high dielectric constant K~ = 80 at 20"C) which interacts with the charged protein molecules to keep them in solution. The addition of less polar soh-ents such as methanol, ethanol and acetone (K j = 21 at 20°C) reduces the dielectric constant of the solution and protein solubility accordingly declines. Fraction­ation by organic soh'ents is sensitive to temperature, pH, ionic strength and the presence of other metal ions. Manipu­lation of these parameters provides flexibility in the selection of separation conditions. Howeycr, org:1l1ic solYCnts haw a tendency to denature proteins and temperatures as low as -lOoe may have to be used to reduce denaturation in a less polar environment.

Probably the most widely used protein precipitation tech­niq ue is salting-out. Addition of salt (either as crystalline

I99

Biotechnology/The Science and the Business

charged and on the mechanism of discharge of solids. All disc-stack machines (Figure 29) contain a set of conical plates or discs separated by flow channels. The thin flow channels mean a reduced "depth" for solids to settle and hence bener performance." The feed is introduced to the rotating bowl by a stationary inlet pipe and passes into the zone between the discs where separation takes place. Under the influence of centrifugal force, the particles traYeI radially outwards until they strike one of the conical discs. The particles then slide down the under side of the disc and are thrown into the space outside the disc stack. The solids may either accumulate here (solids-retaining disc centrifuge), or discharge from rhe bowl continuously (nozzle discharge) or intermittently (solids ejecting with peripheral or axial discharge).

For further details of these de\·ices the literature mentioned in the bibliography should be consulted.

Decanter centrifuge. The decanter bowl discharge centrifuge (Figure 29) is suitable for slurries with high solids contents. Solids are continuously discharged by the helical screw mech­anism. Only reiatiYely low centrifugal forces are feasible.

Filtration

Rotary vacuum (or pressure) filters

Rotary vacuum drum filters appear to be the most commonly employed type of filters in biochemical industry, particularly in antibiotics manufacture and the production of such chem­icals as citric acid by Aspergillus niger fermentations. Rotary filters are available either for suction (or vacuum) operation as in a Buchner funnel or for pressure operation. In the laner the liquid is forced through the filter by the application of pressure on the liquid surface. These filters have the advantage of continuous operation and are useful when sterility and containment requirements are not stringent.

A rotary vacuum filter consists of a drum frame covered with filter cloth (canYas, nylon, Dacron, metal or glass fibre). The internal volume of the drum is divided into radial chambers (Figure 30) to which vacuum may be applied. The

. drum rotates (0.1-2 rpm) partly submerged in an agitated trough of the slurry to be filtered. Application of vacuum (250-500 mm Hg) to the submerged chambers (-30% of filter area) of the drum results in the slurry being drawn through the filter cloth; the initial layer of solids deposit acts as the filter medium. Continuation of suction as the solids­coated drum surface emerges from the slurry bath leads to de"'·a.tering ",·hich may be followed by spray ",·ash. Before the drum re-enters the slurry the solids cake is taken off the filter surface by a knife scraper (doctor blade). Other solids­discharge mechanisms may be used such as strings ",·hich can be lifted off the filter surface. In some cases the filter cloth itself is passed over small diameter rollers to crack the solids cake which then drops off before the cloth returns to the

Spray--.....~ .... heads Cake

recovery Solids ---,r----~

cake

Scraper

,

I Figure 30.

1­drum surface. The type of solids discharge used depends on the properties of the solids. In the cake remonl stage the

Rotary vacuum filter.

suction is discontinued; in some designs the suction chamber may be under positive air pressure to assist cake remon!.

When high filtering capacity and no washing are desired, filter drums with 60 to 70% submerged filter area may be lused. Filtration of fine or gelatinous solids which form imper­meable cakes cannot be carried out effectively with a bare filter cloth which is readily plugged. For such cases a precoat filtration scheme employing filter aids can be used. Filtration of Streptomyces broths often requires filter aids. In this type. of operation a slurry of a filter aid such as diatomaceous earth or cellulose fibres is filtered through the filter cloth to form a porous cake of the filter aid. Subseqciently the broth is filtered through this cake and as a layer of solids deposits on the cake it is scraped off together with a thin layer of the filter aid thereby exposing fresh filtration surface. Suction is maintained throughout the entire cycle to keep the bulk of_~

the filter aid material firmly attached to the filter drum. Precoat filtration is usually limited to cases in which the

solid is not the desired product. Filter aid can be difficult to recover and it may have to be discarded together with the filtered solids. Many biological products are adsorbed to the filter aid material and this loss can be significant particularly when the product is expensive. Laboratory trials are indis­pensible to satisfactory filtration performance for any new application. The broth pretreatment conditions can radically alter its filtration characteristics; changes in pH and tempera­ture, for example, and the length of holding time at these conditions lead to large changes in filtration properties of Streptomyces broths. Proper,~prctreatmentconditions have to

{be. found experimentally. he flux of filtrate, i.e. the volume , . o trate (VL co per unit time (t) per unit filter area (A r), is related to the pressure drop driving force (L'\,P), the

I viscosity 9f the continuous phase (~ld and the flow resistance \.' (RF) by the Hagen-Poiseuille type equation:

I

194

which is usu31ly in aqueous solution (fcrment3tion broth or liquor), when contacted with an immiscible sol\"\.~nt

(cxtractJ.nt), distributes or pJ.rtitions itself between the nvo phases, The extent of p3rtitioning is determined br the par­tition coefficient, k", defined as

(13-t)

'11' b ' For organic acids and bases such as peniCl 111S, enzolc acid, citric acid, erythromycin, kp is strongI)' pH dependent: the salt forms of these compounds (e,g, sod ium benzoate) han' preferential solubility in aqueous media, while the acid

(e,g. benzoic acid) or base (erythromycine) forms show h igher solubilities in less polar, organic soh-ents. Variation of pH is

' . fthus used to alter kp for the desire d extraction. Add Itlon 0

. I f 'd d b dinorg311lc sa ts, atty an s, etergents, etc., can e use to manipulate kp which is also affected by temperarure. Organic phases composed of mixtures of two or more organic solvents have been employed and in these cases k p can be altered by changing the soh-ent composition,

The ratio of the total amount of solute in the two phases is knO'lvn as the "degree of separation" (G) and depends on the "olumes of the soh'em and the broth:

G == C,oken{Vsohen, == k Vsohenr (135) Cbrorh ' V bro,h p V brOIl,

Eq u:ltions (13-t) and (135) are eq uilibri urn rebtionships. Various types of extraction equipment arc a\'ailable and

haw been discussed in chemical engineering literature, In the an tibiotic industry centrifugal extractors such as the Podbidniak extractor arc commonly used for very r3pid ex­traction, Excessi"e exposure of the antibiotic or other product to extraction conditions which may be dcliterious (e.g, low pH in penicillin extraction) is a\'oidcd and because of the high centrifugal fields the form3tion of stable emulsions is reduced. Either whole broth or clear liquor may be extracted but the extraction of \I,hole broth can lead to problems with blockages due to solids. The extractors consist of se\'er31 perforated concentric shells attached to a central sh3ft which acts also as the inlet and outlet for the fluid streams. The sh3it and the shells rotate; the dense liquid is fed to the innermost shell through the shaft and moves radially outward under the action of centrifugal force, A light liquid fed at the peripherr of the outermost shell moves inward counter current to the flow of heav)' liquid, The continuous circular interface (major interface) bet\l'een the two liquids lies somewhere bet,,'ecn the inner and outermost shells and its position can be controlled by Drying the operating conditions. Selection of optimal operating conciitions for any new application re­quires appropriate hboratory trials.

Proteins and other biopolymers which show reason3ble solubility only in aqueous solutions, or are denatured in organic soh'ems, rna)' be purified by liq uid-liquid extraction

Fermenter technology

bct""een (\\'0 aqueous pluses. Immiscible aqueous ph3ses for such applications arc produced by dissoh'ing high concen­tr3tions of twO different polymers or a polrmer-s3lt combi­nJ.tion in ""ater. The reading list should be consulted for additional information on this technique,

Chromatography

Industrial application of chromatography as a separation tech­nique is a recent phenomenon, but already large scale pro­duction chromatography h3s pro\'en itself particularly \'3luable

v

for protein purifications. The production of insulin, hormones and medicinal enzymes makes use of this technique \I·hile

blood plasma fractionation by chromatography is also rapidly developing, Ver)' complex mixtures can be separated or re­

~ soh-ed into their components._--"

Operating principles

All chromatographic sep3r3tions depend on physicochemical interactions between the dissoh-ed components of a mixture and a stationary phase. The latter is usually a solid (or a liquid supported on a solid) contained in a packed column. The mixture is applied to the column 3S' a small volume of solution. The column is then washed or eluted with a soh-ent which constitutes the mobile phase. Different components of thl? mixture move down the column at different rates depending on how strongly a particular component interacts with the stationary phase. The components which associate strongly \I,ith the stationary phase flow down the column at a slower rate than those which do not bind strongly to the solid packing. As a result of their different velocities the different components of the mixture separate as they mo\'e down the column and can be collected as separate components. The phenomenon is analogous to what happens on a racing track: aU contestants arc at the starting position in a single line: part \I'a:-' down the track, however, they have separated and arc at different distances from the starting posi tion and from each other because of their different a\'erage speeds.

In industrial biological separations the mobile phase is alw3)'s a liquid, usually an aqueous solution. Howe,'er, the nature of the interactions between the mixture components and the stationary phase can be \'ery different leading to different types of chromatography: they include affinit)·, ion exchange and hydrophobic chromatography as well as gel filtration. \Vhile all these arc useful in biotechnological separ­ations, ion exchange is widely used for protein purification and affinitr chromatography is an especially powerful tech­nique for biologically active substances. Beginning from the time 3 mixture is applied to a chromatographic column anc the flow of eluting soh'ent is surted, a concentration VS, tim I plot of the components emerging from the column can b· plotted as in Figure 41. The peaks correspond to the tWI components the mixture, A and B. The extent of separatio

20

'--'" Biotechnology/The Science and the Business

(128)

which indicates high density beads and high agitator tip speeds (for higher bead velocity, Ub) for good disruption. The optimum density of the bead material is dependent on the apparent viscosity of the cell slurry inside the mip because the viscous drag tends to reduce the bead velocity. The optimum density of bead material may be estimated using the equation.

Ps = 1016 + 183 ~l.p (129)

for 1 :s; ~l,.p :s; 20 Pa s. Bead diameter and bead loading are other important considerations in a disruption operation. Generally the disruption increases with increasing bead loads and so does the po"·er consumption and the production of heat. Experience shows that bead loading should be 80 to 95% of the void volume of the disruption cbamber. Lower loads lead to poor disruption efficiency.

In general, more rapid disruption is achieved with smaller beads, the optimal bead size being dependent on the size of the microbial cells being disintegrated: the smaller the cell size the smaller should be the bead diameter. For fungal hyphae, for example, bead sizes> 1 mm may be satisfactory. For animal and plant cells and for yeasts the bead size should be < 1 mm. The lower practical limit on bead size is about 0.3 mm and for small bacteria the cell disruption on a single passage through the mill may not produce satisfactory disrup­tion performance. Typically, two or more passes may be needed to cause sufficient disruption of bacterial slurries whereas a single pass may be enough for the larger yeast cells.

Other factors such as the concentration of cells, the location of the desired enzyme within a cell and the strength of the cell wall affect product release by disruption. The cell wall strength depends on the growth environment and the growth stage at which the cells are harvested for disruption. Some of these considerations have been reviewed elsewhere.:!l

The high pressure homogenizer

The high-pressure Manton-Gaulin APV type homogenizer is among the most widely used liquid shear disruption devices. The high pressure homogenizer consists of a positive dis­placement piston pump with one or more plungers. The cell suspension is drawn through a check valve into the pump cylinder and, on the pressure stroke, is forced through an adjust?ble discharge valve (Figure 37) with a restricted orifice. As the cell slurry passes between the valve and the seat its velocity increases rapidly to approximately 290 ms- J with a corresponding decrease in pressure so that cavitation bubbles form. The product velocity decreases again as the suspension leaves the valve seat area causing the bubble to implode. The shock energy released together with the associated turbulence cause the disintegration of cell wall. The impingement of cell

Disrupted cells

Feed __

Handwheel for valve positioning

Figure 37. High pressure homogenizer valve assembly. See reference 21.

slurry on the impact ring (Figure 37) possibly contributes to disruption. Because both cavitation and impingement pro­cesses are velocity associated, the pressure drop across the vah·e (i.e. the difference between the operating and the atmospheric pressure) influences both of them and affects the rate of disruption. The protein released by disruption depends on the pressure difference (tlP) and the number of passes (N ) through the valve as followsp

dR .. -d = kD tlp· (Rm - R) (130)Np

where (R,n R) is the amount of protein remaining to be released. Equation (130) can be integrated for R = 0 at N p = o and R= Rat N p = N p, to

In [Rm/(Rm - R)] = kD VJp tl p. (131)

The exponent a in the pressure term in equation (131) depends both on the microbial cell being disrupted and on its growth history; cells grown on simple synthetic media are generally less robust. For Saccharomyces cerevisiae and Escherichia coli, a values of 2.9 and 2.2, respectively, have been found. Within limits, the disruption process is independent of the concen­tration of cells in suspension. The optimum slurry viscosity and solids concentration ranges tend to be narro""er for the high-pressure homogenizers than for the bead mills. The maximum slurry viscosity should normally not exceed 1 Pas for the homogenizer although more viscous material may be processed in some circumstances. Similarly the maximum acceptable particle size is about 20 ~lm; a lower size (- 2 ~lm)

is preferable. The homogenizer is not suited for fungal broths such ~s those of Aspergilli or for clumps of plant cells. The

I _.

)

~1171t••!I'ij1••• illirmlilii _

In a \Yell packed column the channelling effects (parameter C) arc independent of flow and arc minimized by using small spherical panicles with a narrow size distribution as the packing material. Although theoretical reasoning indicates that a bed of small panicles would give better separation efficiency, the operational requirement of acceptable pressure drop through the bed places lower limits on the size of bed material. The flow rate through a packed bed of rigid particles for a giwn pressure drop is calculated using the Kozeny­Carman equation

_ k, d/ ~p ~__UL - . 7 (142)

L fll (1 - (5)­

which applies for the laminar flow which normally occurs in chromatography columns. This equation is useful because it indicates the effects of bed midage (£s), pressure drop (~P), bed height (L), eluent ,oiscosity (fIr.) and the particle size (d p) on the superficial velocity of the eluting solvent. The packing particle size (d p), the eluent velocity (Ur.) and the molecular diffusivity (Dr.) of the solute in the eluent can be combined into a dimensionless reduced ,-clocity (Vr) given by

v = Uld p (143)

rl

D

For good separation performance Uland dp should be selected to give a reduced velocity in the range 3-10. This criterion, howeycr, yields quite low values of Uland in commercial practice the separation performance is often sacrificed to attain reasonable process throughputs.

Packing material

Selection of the column packing is the most important step in chromatographic separations. Properties of the components of the mixture and the operational requirements determine the most appropriate column packing for the given separa­tion needs. lvlany of the chromatographic packing materials used in bioseparations are porous, hydrophilic substances (cross-linked pol:;acrylamide, agarose, cross-linked dextran, cellulose) which give rise to deformable particles. This restricts both the maximum pressure drop th;lt may be used across the column and the column height to relatively low values to avoid the bed compressibility problems_ Although packing materials with improYCd mechanical properties are becoming available the practical solution to avoiding bed compressibilitY while maintaining satisfactory separation performance is the use of stacked columns. These consiSt of several packed sec­tions, typically no more than 30 cm thick, arranged in series in a stack to give the necessary column length (Figure 43); each section of the stack is a self-comained packed bed. Because the loading capacity (amount of mixture sample that may be handled in a single batch operation) is dependent on the amount of packing material, and because of the restrictions on height of any packed section, the commercial columns

Fermenter technology

Feed

r-----,...,---rl~--Connecting

Staged column

pipes

Product Figure 43. Stacked chromatography column.

tend to have large diameters_ FIO\y distribution, collection and redistribution of flow between packed sections require careful specialist design so that separation achieved is not lost br remixing of components during these transferring oper­ations. Column chromatography as currently practiced is a batch separation technique: a batch of mixture is applied to the column followed by elution "I\-ith a continuous soh-em flow. HO"'ever, column designs which allow continuous ap­plication of mixtures are being deYeloped. ,The literature in the reading list should be consulted for details.

Drying

},!any products of the biochemical industry such as Yaccines, enzymes, pharmaceuticals, etc, ha,'e to be dehydrated for preservation. Dry products keep well and are easy to package and transport. Several types of drying operations are em­ployed; spray and freeze drying are particularly important for thermolabile, biologically active products.

2°3

Biotechnology/The Science and the Business

Spray Drying

Spra)' drying is a method for rapid, continuous, drying of solutions, emulsions and slurries. Pressure or centrifugal atom­izers or gas-liquid jets are used to generate a fine spray of solution droplets which are brought into continuous contact with hot air in a large chamber (Figure 44). Large droplet surface area and small droplet size ensure high evaporation rates so that drying times are but a few seconds. The flow of air is usually cyclonic. The dimensions of the drier must be such that the droplets do not reach the walls until sufficiently dry to prevent sticking and burn-on. A drying chamber tends to be quite Lirge: 1-10 m in diameter being common. The dry powder settles to the bottom from where it is removed either pneumatically or mechanically, or by a combination of these methods.

Advantages of spray driers are: continuous operation, powdered product requiring no further size reduction and rapid drying which leads to good product quality particularly for heat-labile materials but relatively low thermal efficiency is a limitation.

Aseptic spray drying equipment is available. All the air used is filter sterilized and the drying and solids-handling chambers operate under slight positive pressure. The instal­lations can be operated leak-tight and are sterilizable. Anti­biotics such as streptomycine sulfate for direct injection can be spray dried. Highly heat-labile products like some enzymes and blood sera can be successfully spray dried. Microorganisms may be spray-dried for presen'ation and use as SCPo

Air

Drying--. oir

Cyclone

'-----------''-----.. Pro duct

Figure 44. Schematic of a spray drier.

Freeze Drying

freeze drying is the most gentle of the drying methods. The material to be dehydrated is frozen and the ice crystals sub­

2°4

limed by slight warming without thawing. The process may be carried out at atmospheric pressure or under vacuum and sublimation assisted with infrared or microwave radiation, or by contact heating. About 2800 kJ of heat needs to be supplied" for each kg of ice remo\"ed. Freeze-dried products are easy to reconstitute (rehydrate) for use; thus, vaccines, blood plasma, hormones and enzymes are often freeze-dried but a disad­vantage of the technique is the long processing time.

BIOPROCESS CONTROL

In order to ensure optimal functioning of a bioprocessing plant se\'eral processing parameters need to be monitored and controlled. Temperature, pH, product and substrate concen­trations, dissoh-ed oxygen, and material flows are a few of those which may have to be followed over time and manipu­lated in some predetermined way so as to obtain the desired product yields at minimal cost. ,

Computer-based control systems are increasingly encoun-"-,,, : tered in biochemical processing plants and operations such as ' in-place cleaning, filling, sterilization sequences are often fully automated. Control of the biochemical reactor or fermenter is generally limited to control of pH, temperature and dissoh-ed oxygen. A typically instrumented fermenter is shown in figure 45. More extensiw control of fermentation processes is desirable but it is restricted by two main factors: (i) the availability of online sensors to measure the biological and physicochemical parameters needed to follow the progress of fermentation remains limited; and (ii) our limited ability to interpret the available information in the context of the bio­logical system so that the information obtained can be used as a basis for control. Substantial research effort is underway in overcoming these limitations.

Sophisticated control of fermentation systems presuppose' the existence of a mathematical description - or model - 0).-­

Inoculum --+i:~--

Anli!oom --+i:>,<J---{

Acid I Alkc Ii -t¥J----H

I,

Pru~.,I'~ indicatorl eonltollel

r-'----l>'<r-_ Ex hou sf

rlR rIo." indicalor I recordu

~l---Feed

I

~,--L-~_ Cooling woterl

Do~nstream processin9

Figure 45. Typi':'ll fcrmenter instrument.ltion.

steam

f'

ilHiI~IIl.'I'.------------------~

the process. The kinetics of gro,;\·th and prod uct ~ol:mation in combination with the physical system characterlstlcs form a set of equations which constitute the process model. One of the aims of control is optimisation of the process. III batch or fed-batch cultures, profiles of progressive changes ill environ­ment:l! conditions (temperature, pH, dissolwd oxygen, sub­st;ate concentration) arc determined for maximization of product yield. In continuous c~ltivation, op.timisati?n is used to select an environmental regIme for maxImum blOmass or product formation.

A series of chaptc'rs in Comprehensh'e Biotechnology (see further Reading) prm'ide in-depth insight into bioprocess control and control instrumentation.

THE OVERALL PROCESS

6-Aminopenicillanic Acid

1 Penicillins arc among the most widely used antibiotics. About

I 10,000 tonnes of benzyl penicillin (Penicillin-G) arc produced per annum by fermentation with Penicillium chrysogenum.

I Penicillin-G (PEN-G) itself is of limited use and most of it is comerted to 6-aminopenicillanic acid (6-APA) which is a!, precursor for the manufacture of semi-synthetic penicillins

! (Figure 46). Although complete chemical synthesis of 6-APA I is possible, it is economically not feasible on an ind us trial

scale and an enzymetic method is commercially used for the com'ersion of PEN-G to 6-APA. The enzyme responsible for the transformation is penicillin acylase (Ee 3.5.1.11) produced by the bacterium Escherichia coli. Several process varients may be used: (1) E. coli cells grown separately are killed and brought into contact with a solution of PEN-G at 37°C in a batch slurry reactor. After one use, however, the cells lose most of their enzyme activity due to lysis and have to be replaced. This is obviously a drawback. (2) Li\'e, immobilized E. coli cells may be brought in contact with a continuous flow of PEN-G solution. Either packed bed or suspension reactor modes may be used. Side reactions and contamination by substances required to sustain the cells are some possible problems. (3) The enzyme penicillin acylase may be extracted from E. coli, immobilized and used in a continuous reactor. Scheme (l) has certainly been used commercially and the other two routes probably have been, too. The immobilized enzyme route is examined further in the following section.

The overall process schematic flow sheet23 is depicted in Figure 47. A high :'ielding strain of E. coli grown in a batch Stirred tank is used to obtain the enz)'me. The cell slurry is concentrated by centrifugation and the cells disrupted in a ~igh pressure homogenizer. The enzyme penicillin acylase is tsolated by a t'>\'o stage salt precipitation in a low-shear helical ribbon impeller vessel. The first step precipit:ltes the nucleic acids which are removed by centrifugation along with

Fermenter technology ... .. .. _.. _,~ _._-_.~

PenicIllin - G

6 . Aminopenicillanic acid .!­ R .

R-CONH1~~ // C0 2 H

o Semi· synthetic penicillins

Figure 46. Semi-synthetic penicillins from benzylpenicillin (Penicillin-G).

n E. COLI CELLSr FROM FERMENTER

r-------, NUCLEiC

CELL ACID !---.....'CENTR I FUGAT leNDISRUPTION PRECIPITATION

'---~

PEN -G

ENZYME PURIFICATION a IMM08111ZATION

I-c--""'-<

ALKALI

BUFFER ------w 6 - APA TO RECOVERY

Figure 47. Schematic process flow sheet for production of 6­aminopenicillanic acid.

the cell debris. The enzyme precipitate undergoes some further separation to a relatively pure form which is immobilized on amberlite beads. The immobilized catalyst has a half-life of the order of several months under the normal reactor COI1­

2°5

Biotechnology/The Science and the Business

ditions of 37"C and pH 6-8, long enough for industrial

purposes. The kinetics of the deacylation reaction are such that it is

strongly inhibited by 6-APA and is also susceptible to some substrate inhibition. Detailed calculations indicate a packed bed plug flow reactor for the conversion. However, as the deacylation reaction proceeds, the pH of the reaction medium drops and good pH control is necessary to avoid damage to the product, substrate and enzyme catalyst. Control of pH is difficult in packed columns and in this reaction is achieved by the addition of alkali to a buffered reaction medium. Rapid and uniform mixing is necessary throughout the reactor in order to avoid high local pH values which significantly reduce the half-life of the enzyme. To satisfy the operational requirement of good reactor mixing and at the same time to approximate the reaction system to a plug-flow type, a battery of four continuous stirred tank reactors (CSTRs) in series is used for the reaction. The use of four CSTRs is about optimum and a PEN-G conversion of more than 95% is theoretically possible. Addition of a fifth reaction stage in series would increase the conversion only slightly since the productivity of the catalyst in this stage would be low because it would be subject to inhibition by the reaction products which reach their maximum concentration in this last stage. A fifth stage would, of course, increase the capital and operation costs.

The process plant for even a relatively simple process such as the production of 6-APA can be surprisingly complex when in continuous operation. The requirements for automatic cleaning and catalyst replacement as well as the need to accommodate the associated flow schemes leads to extensive

.pipework, valves, pumps and other ancilliaries while the whole plant has, of course, to be built to hygienic standards.

CONCLUSIONS

This overview of bioprocessing and bioprocess engineering has naturally been limited to basic considerations and a selec~

tion only of processing operations has been presented. While very many factors combine to influence the final outcome, it is quite clear that the design of both equipment and operating protocols demands an integrated approach to the process under consideration.

NQMENCLATURE

. Roman Symbols

A Arrhenius parameter (S-I); parameter ineq. (139) l(m2s- )

A b Free area for flow between riser and downcomer (m2

)

Ad Cross sectional area of downcomer (m2)

B b C 6,C C':' Cbro,h

Ccrit

CG CGi

C i

CL

CLi

Cp

Cs

CsoIvcnt

Ct' C'X c D De Dr Dj

DL

d dB de

°dep

dj

dT E 6,E ELr

e G g H

,0 .. i;

-I Filter area (m2

) IHeat transfer area (m2) 1

Arrhenius parameter for nutrient denaturation (S-I)

Surface area of panicle (m2)

Cross sectional area of riser (m2)

Parameter in eq. 63 «m-Is)b) .

Gas-liquid interfacial area per unit dispersion volume (m- I) Gas-liquid interfacial area per unit liquid volume ~m-I)

Parameter in eq. (139) (s) Parameter in eqs 63 and 67 (-) Parameter in eq. 139 (m) Concentration difference (kg m -J) Saturation or equilibrium concentration (kg m -J) Concentration in broth (kg m -J) Critical oxygen concentration (kg m -J) Concentration in bulk gas (kg m -J) Interfacial concentration on gas side (kg m -J) Concentration of species i (kg m-J) Concentration in bulk liquid (kg m -J) Interfacial concentration on liquid side (kg m -J) Specific heat capacity (] kg- I OC- I) Concentration of solids (dry wt./vol.% (g/lOO mL)) Concentration in solvent (kg m- J

)

Tracer concentration (kg m- J)

Equilibrium tracer concentration (kg m -3) Constant (eq. 64, 127) (as appropriate)

lDiffusivity (m2s- )

Effective diffusivity in catalyst (m2s-I) Reciprocal of fraction of unreleased protein (-) Impeller diameter (m)

lDiffusivity in liquid (m2s- ) 0'--../

Diameter (m) Sauter mean bubble diameter (m) Column diameter (m) Connecting pipe diameter (m) Diameter of the ith bubble (m) Tank diameter (m) Total amount of enzyme in reactor (kg) Activation energy (] kmol- I) Liquid phase axial dispersion coefficient in riser (m2s- l

)

Actintion energy for nutrient denaturation (] kmol- I

)

Enzyme concentration (kg m -J) Degree of separation (eq. 135) (-) Gra\Oitational acceleration (ms- 2)

Height equivalent of a theoretical plate (HETP) (m) Heni»s law constant (-) H for component A (m)

206

~T7

~ ~

'1Jvlinimum HETP (m)Hnl Dispersion height (m)hD Heat transfer coefficient (inside film) h, (J S-l m-2°C- 1)

Heat transfer coefficient (inside fouling) 'Jh" (J s- l m- 2°C- I

)

Unaerated liquid height (m) hL Heat transfer coefficient (metal wall) hill

. (] S-l m- 2°C- I )

Heat transfer coefficient (outside film)ho (] s- l m-2°C- 1)

Heat transfer coefficient (outside fouling) hot (] S-lm - 2°C- I )

Ionic strength (-)I FI (k -., -I 3 -., -I)ux 'g m -s or m m -sJ 2O;..:ygen flux (kg m- s- l)J02 Number of CSTRs in series (-)J

K Consistency index (Pa.sn)

K' Constant (eq. 56) (-) Frictional loss coefficient (bottom) (-)K B

Kc Casson viscosity (Pa.s)

KL O\'erall mass transfer coefficient based on liquid film (ms- I) Saturation constant, Michaelis constant (kg m -3)Ks Salting-out constant (eq, 133) (kg m- 3)KSL

KT Frictional loss coefficient (top) (-) Kr Constant (eq. 126) (m- I)

K~ Dielectric constant (-) k Mass transfer coefficient (ms- I) kD Disruption rate constant (S-l) kJ Specific death rate (S-I) kt Filter constant (m- I) kG Gas film mass transfer coefficient (ms- I) k Constant (eq. 54) (-) k

I

L Liquid film mass transfer coefficient (ms- I) k Specific nutrient denaturation rate (S-I)ll

kp Partition coefficient (eq. 134) (-) k r Rate constant (as appropriate) ks Solid-liquid mass transfer coefficient (ms- I) kT Thermal conducti\ity (\'\I m-IOC- I) k, Constant (eq. 142) (-) k , - 3 Constants (eq. 113) (as appropriate) L Length (m) L(p Connecting pipe length (m) . S'! M Mass flow rate of gas (kg m-3) ....... It~ m Mass of bead (kg) N Microbial concentration (m -3 or kg m -3) N Impeller speed (rps) N A Number of theoretical plates for component A

(-) N 0 Initial microbial concentration (m -3) N p N umber of passes (-) n Flow behaviour index (-) n·, Number of bubbles with diameter d i (-)

P ~P

~PFd

~Prr

PG Po ~PTl\1

Q QH

R R R( Re

Rei Rf

J3.~1 Rm

Rp

Rs

R... R" r

rB ri

So Sp T ~T

To T, t

tA.B

t c

t(ool td

t c

tG t g

th.:-.u

tholJ

to top

t r

U

Pow" ("ngm'd~:e,:~:~t:~:::~~;, Pressure drop (Pa) - ­Frictional pressure drop (downcomer) (Pa) Frictional pressure drop (riser) (Pa) Gassed power (W m-~-.,-~ Ee. '"'I (,,)-1 jYJ)

Power number (-r- \"""vJ' <I

Transmembrane pressure (Pa) . lVolume flow rate (m3s- )

Heat transfer rate (J S-I) Specific oxygen consumption rate (kg 02/kg cell· second) Gas constant (J K- I kmol- I) Protein concentration (kg m -3) Resistance due to solids cake (m- I

)

Reynolds number (-) Impeller Reynolds number (-) Total flow resistance (m- I) Resistance of membrane (m -1) Maximum releasable protein (kg m -3) Particle radius (m) Rate of substrate consumption (kg S-I) . Volumetric reaction rate in catalyst (kg S-l m -3) Resolution (-) Radial distance (m) Radius of bowl (m) Initial radial position (m) Resistance of filter medium (m- I) Rate of surface renewal (S-I) Substrate concentration (kg m -3) Schmidt number (-) Sherwood number (eq. 110) (-) Concentration in bulk liquid or concentration in liquid phase at solid/liquid interface (kg m- 3) Initial substrate concentration (kg m -3) Protein solubility (kg m -3) Absolute temperature (K) Temperature difference (OC) Initial temperature (oq Sterilization temperature (OC) Time (s) Retention times of peaks A and B (s) Circulation time (s) Cooling time (s) Mean doubling time (s) Exposure time (s) Gas phase residence time in liquid (s) Mean generation time (s) Heating time (5) Holding time (s) Zero time (s) Operation time (s) Residence time (s) Average flow velocity (ms- t) Bead velocity (ms- I)

2C

I Biotechnology/The Science and the Business

UH

!lL ULc

U Ld U Lm

U Lr

U m .,X

UG U Gr

UT u, Vbroth

Vc VG

"h YL YLd VLr

VpV r /

V.oJ-·m, w

\VIA, WB

X X Xo

X

Xj

Xo

Xc

Z

ZB Zj

Overall heat transfer coefficient (W m - 2oC- I ) ~ll\'(I Fluid viscosity at wall temperature (Pa.s)

Liquid velocity (ms- I ) ~lm :Maximum specific growth rate (S-I)

Mean liquid circulation velocity (ms- I ) !J.N Specific cell number growth rate (S-I)

Superficial liquid velocity in downcomer (ms- I )

U L at H m (ms- I )

~lS

PL Suspension viscosity (Pa.s) Liquid density (kg m­ 3

)

Superficial liquid velocit)· in riser (ms- I ) Ps Density of solids (kg m -3)

Maximum velocity (centreline velocity) (ms- I ) L Sigma factor (eq. 117) (m2

); summation

Superficial gas velocity (ms- 1) o Interfacial tension (N m -I)

Superficial gas velocity based on riser (ms- I ) L Shear stress (Pa)

Tip speed (ms- I)

Terminal settling velocity (ms- I)

Volume of broth (m3)

Chamber volume (m3)

Volume of gas in dispersion (m3)

Liquid "olume (m3) .

I.nterstitialliquid velocity (ms- I)

YL in downcomer (ms- I)

VL in riser (ms- I)

Volume of particle (m3)

Reduced velocity (eq. 143) (ms- I) Volume of soh-ent (m3

)

Dry weight of accumulated biomass (kg) Basal width of peaks A and B (s) Solids concentration (dry) (kg m-3

)

Conversion (-) Initial concentration of solids (dry) (kg m-Mass fraction of oxygen (-) Mass fraction of oxygen (inlet) (-) Mass fraction of oxygen (outlet) (-) Circulation path length (m) Filter depth, axial distance (m) Bowl length (m) Charge on species i (-)

Greek Symbols

ex Parameter (eq. 84) (ms- I) ex Mean specific cake resistance (m kg-I) ~ Parameter (eq. 84) (-); constant (eq. 133)

(kg m­ 3)

y Shear rate (S-I) o Film thickness (m) 0G Gas film thickness (m) 0L Liquid film thickness (m) Os Thickness of solids layer (m) E O\wall gas holdup (-) Ed Downcomer gas holdup (-) Er Riser gas holdup (-) ES Void fraction of solids (-) 11 Effectiveness factor (-) ~l Specific gro'wth rate (S-I) ~l.p Apparent viscosity (Pa.s) !J.L Liquid viscosity (Pa.s)

La Yield stress (Pa) <p Thiele modulus (-) 1jJ Constant kL/d B (eq. 88) (S-I) W Angular velocity

Abbreviations

6-APA 6-Aminopenicillanic acid CMC Carboxymethyl cellulose CSTR Continuous stirred tank reactor DOP Dioctylpathalate HETP Height equivalent of a theoretical plate

l~

I,

, !

. I II,

I I j J

..~

HTST High temperature and short time sterilization PEN-G Penicillin-G (benzylpenicillin) SCP Single cell proteins

3)

REFERENCES

1. Coulson, J.M. and Richardson, J.F. (1977). Chemical Engin­eering, vol. 1 (3rd cd.) Oxford: Pergamon Press.

2. Bailey, J.E. and Ollis, D.F. (1977). Biochemical Engineering Fundamentals, New York: McGraw Hill.

3. Fair, J.R., Lambright, A.J. and Andersen, J.W. (1962). Heat transfer and gas holdup in a sparged contactor. Ind. Eng. Chern. Process Des. Develop., 1,33-36. t

f4. Deckwer, W.-D. (1985). Bubble column reactors. In Biotech nology, vol. 2, p. 445. Ed. by H.-J. Rehm and G. Reed>~l Weinheim: VCH. 1

5. Chisti, M.Y. and Moo-Young, M. (1987). Airlift reactors: t Characteristics, applications and design considerations. Chem. t Eng. Commun., 60, 195-242. ,

6. Mann, R. (1983). Gas-Liquid Contacting in Mixing Vessels, f Rugby: Institution of Chemical Engineers. t Akita, K. and Yoshida, F. (1973). Gas holdup and volumetric· f mass transfer coefficient in bubble columns. Ind. Eng. Chern. t Process. Des. De\·elop., 12, 76-80. l

8. Godbole, S.P., Schumpe, A., Shah, Y.T. and Carr, N.L (1984). } Hydrodynamics and mass transfer in non-Ne'1iJtonian solutions !. in a bubble column. AIChE J., 30, 213-220.

9. Nishikawa, ~1., Kato, M. and Hashimoto, K. (1977). Heat transfer in aerated tower filled with non-Newtonian liquid. Ind. Eng. Chem. Process Des. Develop., 16, 133-137. Chisti, M.Y. and Moo-Young, M. (1988). Gas holdup in pneumatic reuctors. Chem. Eng. J., 38, 149-152. Chisti, M.Y. and 1100-Young, M. (1988). Hydrodynamics and oxygen trunsfer in pneumatic bioreactor devices. Biotechnol. ,-< Bioeng., 31, 487-494.

12. Chisri, M.Y. (1989). Airhft Bioreuctors, London: Elsevier Applied Science.

208

~1I'~lilm.rlin.: _

/

hir, 1.R. (1967). Dcsiglling g,L'-;?:/rze,{ re,letors. Chcl11. F.n~.,--J 3. 7-1 O~ly, 3), p. 67. Dcckwer, \V.-D., 1'\~uycn-Tie!1, h., Schull1pc, A. and ScTpCI11Cn,

14. Y. (1932). Ox)'gcn 1I1.15S tr,'''s}<.'r into "cr"ted CMC so/utiom in a babble column. Biorcchnol. Bioeng., 2-1, 461-481.

15. Schul11!'e, A. and Dcck\\'cl, \\.-D. (1987). Viscous media in to:.:..'cr biorclzC{ors: hydrod)·l1.l?n:·c chdhictcristics and 'rh155 (r...tl1S­fer properties. Bioprocess En~ilh:"ring, 2, 79-94.

16. 1\loo-Young,;\1. and KJ\\'ase, l". (1987). Biorcc,etor d('sign for s!ctrryfcnllcnt,ztion sptcms. III Horizons oi l3io.:hcl11iLcll En~in­eering. p. 281, Ed. b1· S. AibJ. Tokyo: Uniwrsiry of Tokyo Press.

17. 1\kr.:huk,].C and Siegel, i\I.H. (1988). Air-/,ft /C<letors in cbemic,d and biological tec!n;olog1'.]. Chcm. Tech. Biorechno!., -II. 105-J20.

IS. Chisri, }.I.Y., Habrd, B. ;lnd i\!oo-l"oung, }.!. (1988). Liquid circulation in airhft reactors. Chem. Eng. SL·i., 43, 451-457:

19. Chisri, Y. and l\loo-Young. 1\1. (1988). Pri,dwion of liquid circulation velocity in airlift re,lcors ",,·it!; biological medi". ]. Chem. Tech. Biorec'hnol., 42. 211--2J9.

10. F!.lschel, E., \Vandrey, C and hula, l\!.-R. (1983). Ultr,,!,ltr,'tion for tbe separation of bioutalysts. Ad,·. Biochemical Engincerin~, 26, 73-142.

21. Chisri, Y. and Moo-Youn~, 1\1. (1936). Dismption of microbi,z/ cells for intr,'cell"I'/r products. Enzyme \linob. Techno!., 8, J94- 204. I3dl, OJ., Ho,lfc, },[ and Dunnill. P. (19S3). Tb" foml,ttion oj protein precipit,lIL'S and their ce':tnfugal reco~'cry. Ad\'. Bio­chemical Engineering, 26, 1-72.

23. Chisri, M.Y. (1932). A riC:';' plUcess for the prod:{("tion of 6­all1inopcnicilldnic acid fmm PO/icillill-G. The Polytechnic Ib:1d,ln Journal, 1(1), SS·92.

FURTHER READING

I. Reierences 2, 4-6, 12 and 20-22. 2. B:liley,].L (19S0). BlOchemic,z/ ""lction engmecrina and bio­, . I hC!}emlC<1 reClclors. C cm. Eng. Sci., 35, 1854 -1 886.0

3. Blursrrom, LE. (19S5). BlOteefmoIogy: Ferment,Uion and do:.;mst:mm processing. Chem. Eng. (IS february), pp. 126-b3.

4. Brown, D.E. and Ka",ma~h, P.R. (1987). Cross-jlo:.;' septlriltion

Fermenter technology

of cells. Process Biochemistry (August), pp. 96-101. 5. Erickson, R.A. (I %4). Disk st,IC~ cemrifuge) in bio[cc!Jlz()iogv.

Chemid Engineering Progress, 8: (December), pp. 51- 54. 6. hechrcr, A. (Ediror, 19S2). A,h-. in Bioebcmicill Engil!et'ring,

25. Berlin: Springer-Verlag. 7. Karel, S.F., Libicki, S.B. and Roberrson, C.R. (1935). Thc

immobilzz,ltio'l of :';'hole ails: engincering principles. C~em. Eng. Sci., -10, 132J-1354. .

8. Kesha"araz, E., Hoare, j\1. and Dunnill, P. (1937). Biocbe·niml engineering a;pects of cell dis rup tio II . In Sep'lrarions ior Bio­rechnology (~I.s. Verrall and 111.]. Hudson, Eds.). Chichesrer: Ellis Horwood. pp. 62-79.

9. Konecny,]. (1977). T!Jeoretic,zI alld pr,'etic,z/ "speCiS of im­mobilized enz/,mes. In Surn!\' of Progress in ChemisnT, 8 (A.F. Scorr, Ed.). Ne'" York:'AC3demi~ Press: pp. 195-251.

Ie. Kroner, K.H., Nissinen, V. and Ziegler, H. (19S7). Impro:'e'! dynamic fib,aion of miaobiu! suspel15iol1s. Bio/Technology, 5 (Seprember), pp. 921- 926.

11. Mackay, D. and Salusbury, T. (1988). Choosing bt'to:xel1 cen­trifug'1tioll and aossf!o:.;' microfiltrdtiorl. The Chemical Engin­eer, April. pp, 45-50.

12. Moo-Young,]I,I. (EditOr, 1935). Comprehensive Biotec!m%gy, vo!. 2. Oxford: Pergamon Press.

13. Moo-Young,}.l. and Chisri, Y. (19S9). Considerations for designing bioreactors for she.1r-sensiti';e wlture. Bioi Technolog)', 6 (No"ember), pp. 1291-1296.

H. OIJshuc, ].Y. (1939). Fluid mixing in 1989. Chcm. En~. Prog­ress, 85 (l\lay), pp. 33-42.

15. Ripperger, S. (J 93S). Engineering aspects and applications of aossf!o:.;· miaofiltration. Chem. En~. TCL·hno!., 11, 17-2j.

16. S.:hu~erl, K., Lucke, ]. and Oe!s, U. (1977). Buf,blc cu!:"nn biore'ICiors. Ad,·. Biochemical Enginecrin~, 7, J -84.

17. Siegel I, ].H., Dupre, G.D. and Pirkle, Jr., J.e. (19S6). Chromatographic sep,/mtions in a cross-f!o':.;' MSB. Chcm. En~. Progrcss, 82 (Noycmber), pp. 57-61.

1S. Turunjian, R.S. (1985). Scale-Ifp consider.1tions for memf,r,mc processes. Bio/Technology, 3 (jld:·), pp. 615- 626.

19. Webb, e., Black, G ..l\l. and Arkinson, B. (Edirors, 1936). Process Engineering Aspats of Immobilized Cell Systems, Rugby: Instirurion of Chemical Engineers.

20. Whirl', ]l,I.D. and Marcus, D. (1938). Disintegration of miao­organisms. In Adnnces in Biorechnological Processes (A. Mizrahi, Ed.). New York: Alan R. Liss, Inc. pp. 51-96.

2

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