Castilho Leda R., Moraes Angela Maria (Ed.) Animal Cell Technology: From Biopharmaceuticals to Gene Therapy

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uous process can be expressed by multiplying the dilution rate by the

product concentration (DP), it is easy to conclude that low cell concentra-

tions result in low product concentrations and low productivities, making

this process not very attractive for industrial applications.

9.4.4 Continuous cultivation with cell retention (perfusion)

Widely known as perfusion, this operation mode presents the highest

productivities and, at the same time, the highest operational complexity.

Since the innovating work on a spin-ﬁlter in 1969 (Himmelfarb et al.,

1969), the use of this operation mode has become more and more popular,

both at laboratory and industrial scales (Chu and Robinson, 2001).

(B)

Cell concentration (cells/mL)

0.0E 00⫹

8.0E 05⫹

1.6E 06⫹

2.4E 06⫹

3.2E 06⫹

Viable cells

0.6d

⫺1

0.7d

⫺1

0.8d

⫺1

(A)

F, S

F, S, X

10 15 20

Time (days)

Figure 9.16

Continuous cultivation: (A) typical operational conﬁguration; (B) cell

concentration as a function of dilution rate for a continuous culture of a myeloid

transfectoma producing a humanized monoclonal antibody (Center of Molecular

Immunology – CIM, Cuba). The different steady states are indicated by rectangles

and the line has been included to facilitate interpretation of the cell concentration

proﬁle.

242 Animal Cell Technology

In this operation mode, it is possible to mitigate the major limitation of

continuous cultures, that is, the low productivity due to the loss of cells in

the bioreactor outlet. In perfusion, this issue is overcome by using a cell

retention device to maintain cells inside the bioreactor. Figure 9.17 shows

a scheme of a stirred-tank bioreactor operating in perfusion mode, as well

as the kinetic behavior of a perfusion run.

Equations 8 and 9, previously shown for the substrate and product mass

balances in continuous cultures, can also be applied to this cultivation

mode. However, due to cell retention, the equation that expresses the cell

mass balance is altered, when compared with the mass balance of cells in a

continuous culture (Equation 11):

0 1020304050

(A)

F, S

Cell retention

device

F, S, Xvα

(B)

Viable cells Product

Time (days)

Product concentration (cells/mL)

Product concentration (mg/L)

0.0E 00⫹

1.0E 07⫹

2.0E 07⫹

3.0E 07⫹

4.0E 07⫹

100

120

Figure 9.17

Perfusion culture: (A) t ypical operational conﬁguration, using an external cell

retention device; (B) concentration of viable cells and product in a perfusion

bioreactor culture of a myelo id t ransfectoma producing a humanized mono clonal

antibody (Center of Molecular Immunology – CIM, Cu ba).

Bioreac tors for animal cells 243

¼  X

 ÆDX

(11)

In Equation 11, Æ is the cell passage factor, deﬁned as the ratio between

cell concentration in the perfusate and in the bioreactor (Equation 12).

Æ ¼

v,PERFUSATE

(12)

Cell retention may be total or partial. A total cell retention would result

in Æ ¼ 0, whereas a value of 1 would mean no retention of cells.

Wash-out conditions in perfusion cultures are different due to the

retention of cells. In the same way as Equation 10 was obtained, Equation

13 can be derived from Equation 11:



MAX

¼ Æ D

washout

(13)

For low Æ-values (between 0 and 0.2), the dilution rate may be increased

up to values that are much higher than the maximum speciﬁc growth rate,

without establishing a wash-out condition. The increase in the dilution

rate that can be employed results in a greater availability of nutrients and,

consequently, in an increase of cell and product concentrations. Thus, the

perfusion mode allows a high-cell-density culture to be maintained for a

long period, at a high perfusate ﬂow rate and an elevated product

concentration. This results in productivities that are 1–2 orders of magni-

tude higher than those obtained in continuous processes. In fact, the

performance of perfusion cultures is superior to that of all other operation

modes (Bibila and Robinson, 1995).

The operation of cultures in perfusion mode is possible for almost all

existing bioreactor types. Heterogeneous bioreactors are usually operated

in perfusion mode, and homogeneous bioreactors can be if a solid–liquid

separation device (cell retention device) is used (see Chapter 11).

As indicated for other modes, it is necessary to maintain the inoculum

concentration in the range of 0.1 3 10

to 0.4 3 10

cells mL

–1

,to

minimize the adaptation lag phase at the beginning of a cultivation process.

The proﬁles for cell growth and product formation will depend on the

feeding strategy adopted, on the cell line characteristics, and on the

performance of the cell retention device.

In the case of homogeneous bioreactors, the maximum cell concentra-

tion in perfusion cultures can attain 10

–10

cells mL

–1

. When this opera-

tion mode is used in heterogeneous bioreactors, cell concentration in the

cell compartment can approach the packing limit of tissues, which is in the

order of 10

cells mL

–1

. Product concentrations reported for these pro-

cesses vary considerably, but are most commonly in the range of 100–

500 mg L

–1

. Culture duration can be in the range of several days up to

several months (Bo

deker, 1994).

The feeding strategy adopted in a perfusion process is of great impor-

tance. The ideal approach, that is, the one that results in the highest

productivity, depends on the relationship between growth kinetics and

product formation kinetics. For cell lines that present growth-associated

production, it is recommended to feed culture medium at an increasing

dilution rate to allow maximum cell growth. However, if product forma-

244 Animal Cell Technology

tion by the cells is not associated or inversely associated with growth

kinetics, then the dilution rate should be increased up to a given value and

then remain constant, resulting in a steady state of the main process

parameters (Figueredo, 2002).

An example of the ﬁrst strategy is the so-called cell-speciﬁc perfusion

rate (Ozturk, 1996), which allows maintenance of cells in a constant

environment by exponentially increasing the feed rate according to the cell

concentration proﬁle (Chico and Ja

ger, 2000). A negative aspect of this

strategy is that very high cell concentrations are achieved rapidly, and this

may negatively affect the cell separation device (Deo et al., 1996), and thus

limit the culture duration. This negative aspect can be minimized by

establishing a cell bleed stream for a controlled removal of cells from the

bioreactor. This allows maintenance of a high cell viability, combined with

a stable and high cell concentration, at a level that is compatible with good

performance of the cell separation device. Furthermore, the high dilution

rates used in this case result in short residence times of the product inside

the bioreactor, which can be relevant in the case of products that are very

unstable or that can be easily degraded under culture conditions.

Another feeding strategy, which is sometimes called ‘‘stationary strat-

egy’’, aims at extending the culture duration for months (Bo

deker, 1994).

However, the use of this type of strategy results in a considerable decrease

in cell growth rate (Figueredo, 2002) and in an accumulation of dead cells,

causing the release of cell debris and substances potentially dangerous to

product integrity. To diminish this problem, cell purges are periodically

carried out in order to maintain a more constant and controlled cell

environment (Kempken et al., 1991).

Although in most cases perfusion operation presents several economic

advantages, its industrial use is not as widespread as that of fed-batch

cultures. This is due mainly to false ideas of the technological and

operational difﬁculties and an underestimation of the productivity gains of

this operation mode. A critical evaluation of the characteristics of perfu-

sion operation indicates that the disadvantages have usually been empha-

sized or even exaggerated (Kadouri and Spier, 1997).

The design of bioreactors for perfusion operation is more sophisticated,

which makes the equipment more expensive. However, the productivity

increases obtained by perfusion operation allow the use of much more

compact systems than those operated under batch or fed-batch mode. In

this way, perfusion bioreactors can be up to 10-fold smaller for a given

production scale (Bibila and Robinson, 1995), decreasing the costs not

only of the bioreactors themselves, but also of storage tanks and down-

stream processing equipment.

As regards validation, perfusion bioreactors, which are more complex

and of more recent use, have been classiﬁed in the past as difﬁcult to

validate (Kadouri and Spier, 1997). The longer cultivation runs result in

longer validation processes when compared with discontinuous processes.

However, the growing number of approved biopharmaceutical production

processes based on perfusion mode suggests a change in attitude by the

regulatory agencies (Chu and Robinson, 2001).

In summary, the main technological features of this operation mode are:

Bioreac tors for animal cells 245

(i) more complex operation than in the case of other operation modes;

(ii) higher contamination risk, since it is an open system that operates

continuously for long periods of time;

(iii) higher volumetric productivity;

(iv) short residence time of product in the bioreactor (ideal for the

production of labile molecules);

(v) process optimization basically through manipulation of medium

composition, feeding strategy, and controlled cell removal (cell bleed-

ing);

(vi) scales generally up to 2000 L.

9.5 Aeration and agitation

Aeration and agitation are two important operations in animal cell culture.

Oxygen has a low solubility in aqueous media, with a saturation concen-

tration of approximately 7 mg L

–1

at 378C. This implies that the oxygen

must be provided continuously to the cultures, to ensure that the dissolved

oxygen levels in the culture medium remain at an adequate level. The mass

balance for oxygen in the liquid phase can be written as:

dDO

ðÞ

¼ k

aDO



 DO

ðÞ

 q

(14)

Where:



¼ concentration of dissolved oxygen at saturation (in equilibrium

with the gas phase);

DO ¼ dissolved oxygen concentration;

¼ global oxygen transfer coefﬁcient;

a ¼ gas–liquid interfacial area per reactor volume;

¼ oxygen speciﬁc consumption rate.

This mass balance concerns the liquid phase, since oxygen must be

dissolved in order to be used by the cells. Due to the difﬁculty in measur-

ing the interfacial area (a), especially when oxygenation is carried out by

bubble aeration, it is common to use the product of k

times a (k

a),

known as the volumetric oxygen transfer coefﬁcient, as the relevant

parameter.

The dissolved oxygen concentration at equilibrium follows Henry’s

Law (Equation 15), with the constant H

being a function of tempera-

ture. Since the air contains 21% of oxygen, the partial pressure P

when

using pressurized air is ﬁvefold lower than when using pure oxygen. Thus,

DO* is higher for pure oxygen injection, resulting in a larger oxygen

transfer rate to the liquid medium.



¼ H

 P

(15)

Where:

¼ oxygen partial pressure;

¼ Henry’s constant.

Different studies have established an optimal dissolved oxygen concen-

tration for animal cell culture in the range of 20–50% of air saturation

246 Animal Cell Technology

(Butler, 2004). Too high oxygen concentrations lead to the formation of

free radicals in the medium, which may cause oxidative damage to cells.

The DO concentrations are monitored with previously calibrated electro-

des, whereby the 100% value is adjusted after saturation of culture

medium with air, at the cultivation temperature (usually 378C). The

oxygen transfer system must be designed to meet the demands of the

culture, avoiding situations of limitation or excess of dissolved oxygen.

This is accomplished by supplying air, pure oxygen, or mixtures of both

gases to the culture. Another aim of aeration is to allow dissolved carbon

dioxide originating from cell respiration to be removed in the gas exhaust

stream, to avoid toxic levels (Gray et al., 1996). The removal of carbon

dioxide is known as ventilation.

Different aeration methods have been employed in animal cell culture:

(i) surface aeration;

(ii) aeration through membrane devices or gas-permeable tubing;

(iii) bubble aeration.

The ﬁrst two methods are widely used just for laboratory-scale applica-

tions, since there are limitations for their use at larger scales. In surface

aeration, the oxygen transfer rate is related to the liquid surface area (gas–

liquid interface). When the culture scale is increased, maintaining constant

geometric proportions, the volume increases with the third power of a

characteristic linear dimension of the system, whereas the surface area

increases with its square.

A recent option is that found in wave bioreactors, where the generation

of waves increases oxygen transfer by augmenting both the interfacial area

(a) and the global transfer coefﬁcient k

(Singh, 1999). Therefore, these

reactors are already available commercially at volumes up to 500 L.

The second method uses devices that can be positioned either inside the

bioreactor or in external recirculation loops. This method allows aeration

in the absence of bubbles, minimizing the negative effects that these can

exert on the cells, as will be discussed later in this chapter (Lehmann et al.,

1987). The aeration devices can be based either on membranes or on non-

porous tubes made of oxygen-permeable materials, such as silicon, or on

porous devices made of materials that are impermeable to oxygen, such as

polypropylene. This method also has limitations regarding scale-up, and

its use has been reported just for bioreactors up to 150 L in volume

(Lehmann et al., 1987). Its use at a larger scale becomes unfeasible, due to

the operational difﬁculties associated with assembling, cleaning, and

sterilizing systems that contain hundreds of meters of tubes or membranes.

In heterogeneous bioreactors, external aeration devices are most com-

monly employed, since it is easy to recirculate a cell-free or almost cell-

free stream through an independent oxygenation system. Due to the low

solubility of oxygen in aqueous media, the recirculation speed usually

needs to be very high, reaching up to 100 vvd (recirculation volume per

reactor volume per day).

The aeration method most widely used in animal cell cultures is bubble

aeration, just as in microbial fermentations. This method is simple and

consists of bubbling a gas stream directly into the culture medium, using a

Bioreactors for animal cells 247

sinter sparger, oriﬁce sparger or jet-ﬂow device. Sinter spargers are made

of porous stainless steel and can generate bubbles with diameters in the

order of hundreds of micrometers. Oriﬁce-type and jet-ﬂow spargers

generate bubbles 1 mm in diameter or larger (Puleo et al., 2004). In the

bioreactors, the spargers are generally positioned below the impellers, to

promote a homogeneous distribution of bubbles inside the culture vessel.

Although simple, the use of bubble aeration in animal cell cultivation

requires the solution of certain technological challenges for its implemen-

tation, since one of the most characteristic features of animal cells is their

low resistance to mechanical stresses. Cell membranes have a mechanical

resistance that is much lower than that of microbial cell walls. According

to Castilho and Anspach (2003), shear stresses in the order of 60 Pa are

critical for cell death.

It is reported that bubble disengagement and bursting at the liquid

surface is a major cause of cell death in animal cell cultures (Kunas and

Papoutsakis, 1990; Butler, 2004). The cell membrane is slightly hydropho-

bic under culture conditions, and therefore cells tend to adhere to bubbles

ascending inside a bioreactor. When these reach the liquid surface, the ﬁlm

of ﬂuid that surrounds the bubble begins to be drained by gravity,

decreasing in thickness. The bubble bursts at the moment the mechanical

resistance of this ﬁlm is not sufﬁcient to resist the pressure inside the

bubble. The explosion of a bubble causes large velocity gradients in the

liquid below the bubble. Figure 9.18 shows that the presence of a bubble

creates a cavity in the liquid and, at the moment of bubble bursting, this

cavity is quickly ﬁlled with liquid, generating high velocity gradients

within short distances, resulting in shear stress levels that are sufﬁciently

high to destroy the cells found in this region (Meier et al., 1999).

In general, to minimize damage to the cells, very small bubbles should

be avoided in animal cell cultures, since the interfacial area of the bubbles

Cavity created by

the bubble

Retroceding

film

Liquid

Flow below the

bubble cavity

Air

Interface

Figure 9.18

Dynamics of bubble explosion (adapted from Wu and Goosen, 1995).

248 Animal Cell Technology

is intrinsically related to the kinetics of cell death in bioreactors (Tramper

et al., 1988; Wu and Goosen, 1995). Small bubbles provide a greater total

interfacial area than large bubbles. Therefore, small bubbles are able to

carry more cells to the top gas–liquid interface, where cell damage occurs

due to bubble explosion. On the other hand, small bubbles can transfer

oxygen more efﬁciently, and this allows the use of much lower aeration

rates. Thus, all these factors should be evaluated when determining the

most adequate operational conditions.

The addition of surfactants allows a modiﬁcation of the kinetics of non-

speciﬁc cell adhesion to bubbles. When substances such as methyl cellu-

lose or Pluronic

F68 are added to the culture, the time needed for cell

adhesion to occur is increased (Meier et al., 1999). In this way, the number

of cells adhered to a bubble at the moment of its explosion is lower by

several orders of magnitude. The non-ionic surfactant Pluronic

F68 is so

far the best option, since it efﬁciently protects cells from bubble damage

without signiﬁcantly affecting oxygen transfer. However, its presence may

be undesirable for certain stages of protein puriﬁcation.

Foam formation is another problem that occurs in bubble-aerated

bioreactors, especially in the presence of serum-containing culture media

or at high protein concentrations (Butler, 2004). An uncontrolled accumu-

lation of foam in the upper part of the equipment can occur, resulting

eventually in severe problems due to blockage of exhaust gas ﬁlters. As a

consequence, gas transfer through the surface is seriously affected. To

decrease the negative effects of foam, different approaches can be adopted:

(a) chemical antifoams; (b) foam traps; (c) bubble-free aeration; or (d) low

aeration rates using pure oxygen.

The use of silicon-based antifoams is common in industry (van Bonarius

et al., 1993). However, they should be used with care, since these

substances can be toxic to the cell above certain concentrations. Further-

more, chemical antifoams can pose problems for the chromatographic

puriﬁcation of the product. Foam traps, which are devices mounted in the

upper part of bioreactors to break the foam, have been used successfully at

small and intermediate scales, but are not widely used on a large scale. On

the other hand, low aeration rates using pure oxygen effectively lead to a

signiﬁcant decrease or even complete elimination of foam, but may result

in CO

accumulation in the medium, which is harmful to the cells (Gray

et al., 1996).

Both aeration and ventilation are enhanced by mechanical agitation of

the cell suspension. However, agitation also has other functions in a

bioreactor: to maintain cells in suspension (in the case of homogeneous

bioreactors or suspended microcarriers), as well as to homogenize the

ﬂuid. This homogenization is important to avoid the appearance of dead

zones and of nutrient, metabolite, and temperature gradients, as well as to

promote the transfer of heat and of the different chemical species, includ-

ing oxygen, in order not to limit the performance of the biological system.

When stirred-tank bioreactors started to be used for animal cell

cultivation, many problems related to deleterious effects of agitation on

cell viability were observed. However, it was noted that the use of large

impellers rotating at low speeds could minimize mechanical damage to

the cells. The most widely used impeller types are marine (Chisti, 1993)

Bioreac tors for animal cells 249

and inclined-blade impellers (Michaels et al., 1996). On a small scale,

stirrers with oscillating membranes (Lehmann et al., 1987) have also

been proposed. However, these and other similar systems have limited

efﬁciency and are difﬁcult to scale-up. Currently, marine propellers

continue to be the most widely used impeller type, but also radial-ﬂow

impellers, such as turbines, or combinations of different types are

utilized (Krahe, 2003).

The ﬂow pattern in a stirred bioreactor with bafﬂes is extremely

complex and, generally, turbulent under normal operational conditions

(Brucato et al., 1998). Higher shear rates and, consequently, higher energy

dissipation rates occur near smaller turbulent eddies, that rotate at higher

velocities. The smallest size an eddy can reach under given ﬂow conditions

is known as microscale of turbulence and can be estimated through

Kolmogorov’s equation (Joshi et al., 1996):

º ¼







0:25

(18)

Where:

º ¼ Kolmogorov eddy length (microscale of turbulence);

 ¼ kinematic viscosity;

 ¼ energy dissipation per unit mass.

Several authors have proposed that, when the microscale of turbulence

is of the same order or smaller than the cell diameter, the energy dissipa-

tion due to the eddy action on the cell surface can destroy the cells (Kunas

and Papoutsakis, 1990). In the case of larger eddies, cells can be accom-

modated in their interior, moving at the same velocity as the eddies. Thus,

no velocity gradients are formed that could damage the cells. However,

other authors state that adopting Kolmogorov’s theory to explain cell

damage is inadequate, mainly due to the evidence that the dominating

deleterious effects on cells are due to bubble explosion at the gas–liquid

interface (Kioukia et al., 1992). Furthermore, it has been demonstrated

that in the absence of a gas phase inside the bioreactor, cells are more

resistant to high rotation speeds than when bubbles are present (Kunas

and Papoutsakis, 1990).

9.6 Scale-up

To meet the demand for a product in the market, two approaches are

possible; process intensiﬁcation by increasing process productivity, or

increasing the production scale. The latter approach may be achieved by

increasing the number of bioreactors of a given size or by using bioreac-

tors of larger volume. From the economic viewpoint, generally the second

option is more adequate, since equipment costs usually increase with scale

to the power of 0.6 (Rouf et al., 2000). Beyond that, operation and

maintenance costs increase with the number of pieces of equipment,

indicating that for scale-up an increase in the bioreactor volume is more

advantageous than using several smaller units.

The aim of scaling-up is to reproduce in production scale the conditions

optimized at laboratory and/or pilot scale. Most of the general principles

250 Animal Cell Technology

adopted for the scale-up of microbial fermentation processes are applicable

to animal cell culture. Usually the factors that can inﬂuence volumetric

productivity are investigated when changing the scale in order to identify

those that are most critical. Once they are identiﬁed, similar criteria for

these factors are established for the different scales. The most commonly

used criteria concern the peripheral rotation velocity of the impeller and

the gas ﬂow rate, in the case of bubble-aeration processes.

Before scaling-up a process established at a small scale, it is important

that it is appropriately characterized and optimized. First, a cell line with

desirable production, growth, and genetic stability should be selected. This

selection is usually carried out in multiwell plates and stationary ﬂasks

(Castillo et al., 2005). An adaptation to suspension growth is usually

desirable. Then, the master and working cell banks are prepared, to

cryopreserve the selected cell line. Subsequently, process development is

initiated. Studies on the behavior of the cell line are carried out in

laboratory bioreactors that should have the same conﬁguration as those to

be used at a production scale. Under these conditions, the kinetic charac-

teristics, production pattern, oxygen requirements, and shear resistance of

the cell line are determined (Vı´ctores, 2005). After that, a mode of

operation should be chosen and the operational variables that maximize

volumetric productivity should be determined. Finally, studies aimed at

process scale-up should then be carried out.

The stirred-tank bioreactor is an example of a reactor that is scaled-up

by direct increase of its volume. One of the critical parameters that should

be evaluated at different scales is the rotation speed of the impeller.

Different criteria can be employed:

(i) to maintain constant the power per ﬂuid volume dissipated by the

impeller;

(ii) to keep constant the peripheral velocity of the impeller;

(iii) to maintain constant the volumetric oxygen transfer coefﬁcient (k

a).

Each of these criteria will give different results, so that they should be

carefully evaluated, preferably from experimental results. For the aeration

parameters, it is important to evaluate which is the most relevant phenom-

enon; the capacity of bubbles for oxygen transfer or their negative effect

on the cells (Chisti, 2000). At very large scales, the ventilation is also a

very important issue, mainly if low aeration rates with pure oxygen are

used (Gray et al., 1996).

Since in animal cell culture processes the effects of mechanical stress are

much more relevant than in microbial fermentations (Chisti, 1993), it is

quite common to adopt scale-up criteria that are associated with cell

damage (Joshi et al., 1996), such as constant peripheral impeller velocity,

constant aeration rate, and constant integrated shear stress (Croughan et

al., 1987).

When bioreactors coupled to cell retention devices are used, it is also

necessary to evaluate the scale-up of the cell separation equipment. In the

case of the spin-ﬁlter (see Chapter 11), parameters such as ﬁlter rotation

velocity and the ratio of ﬁltration area to bioreactor working volume are

particularly relevant (Deo et al., 1996).

Bioreac tors for animal cells 251