Professor Wei-Shou Hu Spring 2007 ChEn 5751

  Professor Wei-Shou Hu Spring 2007 ChEn 5751

  Cell Culture Bioreactors Basic Types of Bioreactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .1 Segregated Bioreactors (Dead Zone Present)\Compartmentalized Bioreactors . 4

  Homogenous Reactor vs . Heterogeneous Reactor . . . . . . . . . . . . . . . . . . . . . . . . . . 4 Batch and Continuous Processes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 4 The Operating Mode of Reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 Batch Cultures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 Fedbatch Cultures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 Continuous Cultures . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7

  Material Balance on Bioreactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .8 Material Balance Equation for Reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 8 Tissue culture and disposable cell culture systems . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .10 Microcarriers . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 2

  Cell Aggregates . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 Microsphere Induce Cell Aggregates . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 5 Agarose . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 6 Microencapsulation . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 6

  Cell Culture Bioreactors. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .17 Simple Stirred Tank Bioreactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 7 Airlift Bioreactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 8 Spin Filter Stirred Tank . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 9 Vibromixer . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 20 Fluidized Bed Bioreactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 2

  Basic Types of Bioreactors

  Mammalian cell bioreactors are generally categorized similarly to chemical reactors according to their mixing characteristics. It is instructive to review two ideal reactors: well-mixed stirred tank and plug-flow (tubular) reactor. In an ideal well-mixed bioreactor, the mixing is assumed to be intense enough that the fluid is homogeneous through the reactor. The mathematical description of ideal continuous flow stirred tank reactor is described by the following first-order differential equation.

  d VC ( )

  = F C − F C r V

• A

i Ai o Ao A

  dt V is the culture volume in the bioreactor, CA the concentration of nutrient or product A, t is time, F is the flow rate and rA is the volumetric consumption rate of nutrient or production rate of product A. In an ideal stirred tank reactor, there is no flow bypass and no shunt of substrate from inlet to outlet, no dead zones or clumps of undissolved solid substrate floating around. The addition of a substrate through feeding is instantaneously distributed throughout the entire reactor, and when gas sparging is employed the agitator provides an intimately mixed gas-liquid. It also follows from this assumption that th e stream exiting the reactor will have the same composition as the well mixed fluid in the reactor. The basic model for the tubular reactor (such as hollow fiber and ceramic systems to be described later in this chapter) specifies that the liquid phase moves as a plug-flow, meaning that there is no variation of axial velocity over the cross section. The mass balance for component A in a volume element S z ∂ that described an ideal plug-flow reactor is the following:

  ∂ C ∂ C ^{A A} = − v r

_{Z A}

+ t z

  ∂ ∂

  where vz is the linear velocity in the z direction along the flow and S is the cross-sectional area. Note that we assume there is no liquid dispersion or back mixing. All elements in the fluid move at the same velocity. At steady state (i.e., cell concentration and cellular activities at a given position are not changing with time), the equation becomes

  ∂ c r ⋅ S _{A A} = ∂ z F

  which describes changes of concentration of A along the direction of fluid flow. It is clear that the nutrient concentration will decrease from inlet to the distal end of the reactor, while metabolite concentration increases. The length of the reactor is limited because eventually nutrient depletion or metabolite accumulation inhibits growth and metabolism. These ideal cases of completely mixed tanks or plug flow tubular reactors are situations that can be approximated in small-scale laboratory conditions. The conditions in larger scale process reactors deviate significantly from these ideal conditions. In a well-mixed bioreactor, there are no concentration gradients in either the gas or the liquid phase. In other words, none of the chemical species or cells is segregated in the reactor. The other extreme of mixing is total segregation where there is no interaction between different volume elements in the bioreactor. An ideal plug flow reactor is assumed to be under conditions of total segregation.

  Most bioreactor systems have a mixing pattern between the two extremes and are under partially segregated conditions. In general, laboratory and small pilot plant bioreactors are used for process development and optimization. The fluid mixing characteristics are rather sensitive to the scale of the reactor. Furthermore, plug-flow bioreactors are intrinsically more difficult to scale up than mixing vessels, as the concentration gradient of essential nutrients, oxygen in particular, will inevitably become limiting in the downstream region of the reactor. In considering the selection of bioreactors for mammalian cell cultures, the mixing characteristics and their relationship to scale-up have to be kept in mind.

  Stirred Tank (Well Mixed) vs. Tubular Reactor (Plug Flow)

  The distinction between a well mixed continuous stirred tank reactor (CSTR) and plug flow reactor (PFR) is best illustrated by comparison of their response in the outlet to a step change in feed concentration, consider a continuous reactor that has an inlet stream (feed) and an outlet stream that are equal in volumetric flow rate, the volume of the reactor is thus constant. For the case that the feed stream is colorless, but at time 0 the stream is changed to a feed with red color at a concentration of C If the reactor is well mixed as in a CSTR, as soon _O as the feed stream is switched, the color will be seen immediately in the effluent stream, since the color is distributed instantaneously everywhere including the fluid that is taken out in the effluent stream. The plots shown are the colors seen at the outlet. The red dye concentrate will increase gradually. If the reactor volume is V, it will take longer than the time needed to flow through one reactor volume to reach the same concentration as in the feed, since the dye is also being taken out from the reactor from the beginning. In fact, by solving the differentiation equation, t can be shown that it takes _I about three holding times (3t ) to reach almost the same concentration as in the feed. Now examine the case of PFR. According to the model of PFR, the red color dye will move downstream like a sharp band, since there is no backmixing or diffusion to blur the sharp boundary between the color and colorless streams. So the detector at the exit will detect no color right after the switch to dye solution in the feed. It will not see any color until the “front” of the color feed solution reaches the outlet. The time it will take will be a “hold time”, the exact time of flow a reactor volume into the reactor to displace all the original clear solution in the reactor. As soon as the color comes out in the outlet, the concentration will be equal as in the feed.

  If the reactor is not idealized, obviously the pattern of the dye concentration will be different. For a tubular reactor, the “front” may not be as sharp, rather the appearance in the exit will be more gradual. Similarly in a stirred tank, there will be deviation to the perfect mixing curve. In more severe cases, a reactor may be compartmentalized or segregated. Some channeling may occur to have some portion of the feed stream passing right through, or some portions of the reactor hardly see any feed stream.

  Segregated Bioreactors (Dead Zone Present)\ Compartmentalized Bioreactors

  Concentrations in different compartments may be different Most reactors are not ideally all mixed or plug-flow; segregated zone is not a completely dead zone

  Implication When Growth or Reaction Occurs in the Reactor

  Flow and mixing behavior may have a profound effect on the

  Reactor reaction or growth. Consider a nutrient stream entering the reactor.

  If the reactor is a PFR, the cells in the upstream will have abundant nutrient. As the fluid moves downstream more nutrients get consumed and their concentration decreases. The cells downstream may not have enough nutrients or face starvation. One way to solve the problem is of course to increase the supply rate by using a higher nutrient concentration in the feed or by operating at a higher flow rate. But there are limits on both nutrient concentration and flow rate. Eventually the size of the reactor will be restricted. In a CSTR model, all cells in the reactor see the same environment. The nutrients that feed into the reactor will be distributed uniformly everywhere, either all have abundant or suboptimal levels.

  Homogenous Reactor vs. Heterogeneous Reactor

  Heterogeneous reactor—with a solid phase, e.g., microcarriers in stirred tank, tubular reactor packed with foam. A typical tissue has a cell concentration of about 5X10 /ml. Unless a rector have a very _{8 7} high cell concentration in the middle of 10

  /ml, cell mass is only a small fraction of the culture volume. So, even though almost all cell culture reactors have all three phases, liquid medium, gas bubbles and cell mass, they are often treated as homogenous bioreactors. On the other hand, in addition to high cell density culture there are cases where the bioreactor must be treated as heterogeneous. The solid phase constitutes a large fraction of the culture volume. An examples is the microcarrier culture. Microcarrier beads often constitute 10-30% of the culture volume. In such cases even cell ₇ per concentration needs to be well defined, for example, whether 10 milliliter is referring to total culture volume or liquid volume needs to be specified.

  Operating Mode of Bioreactors Batch and Continuous Processes

  A reactor is called continuous when the feed and product streams are continuously being fed and withdrawn from the system. In principle, a reactor can have a continuous recirculating flow, but no continuous feeding of nutrient or product harvest; it is still a batch reactor. A fed-batch bioreactor usually has intermittent feed. It may or may not have medium withdrawal during the run.

  Example: For instance, Yeast cells (saccharomyces cereviciae)

  can metabolize glucose either to ethanol, or to oxidize it to carbon dioxide, mammalian cells can convert glucose mostly to lactate, or oxidize it to carbon dioxide. Cells in two such types of metabolism are in two different metabolic states. The two metabolic states are characterized by different specific glucose consumption rates, lactate or ethanol production and the yield coefficient for biomass, i.e. different stoichiometric ratio.

  Example:

  For instance, a 1 l culture has 0.3 of solid microcarriers ₉ cells in it. The cell concentration and 0.7 l of medium, with 10 ₉ cells/L-medium. If the glucose is 109 cells/L-culture or 1.43 x 10 concentration in the culture medium decreases from 2.10 g/L (medium) over one day, then the specific glucose consumption rate ₉ is (2.10-1.90) g/L-medium ÷ (1.43 x 10 cells/L-medium) = 1.40 x ₁₀

  10- g/cell-hr. The specific rate calculated would have been very different if one concentration is based on liquid volume and the other is based on total culture volume.

  The Operating Mode of Reactors Batch Cultures

  Batch processes are simple and are widely used, especially in the vaccine industry and in pre-production scales of rDNA protein production. Fedbatch processes are widely used in multi-purpose, multi-product facilities because of their simplicity, scalability, and flexibility. A variety of fedbatch operations, ranging from very simple to highly complex and automated, are seen in current production facilities.

  Fedbatch Cultures Intermittent Harvest

  In general, fedbatch processes do not deviate significantly from batch cultures. For both intermittent-harvest and traditional fedbatch cultures, cells are inoculated at a lower viable cell density in a medium that is usually very similar in composition to a typical batch medium. Cells are allowed to grow exponentially with essentially no external manipulation until nutrients are somewhat depleted and cells are approaching the stationary growth phase. At this point, for an intermittent-harvest fedbatch process, a portion of the cells and product are harvested, and the removed culture fluid is replenished with fresh medium. This process is repeated several times. This simple strategy is commonplace for the production of viral vaccines produced by persistent infection, as it allows for an extended production period. It is also used in roller bottle processes with adherent cells.

  Fedbatch

  For production of recombinant proteins and antibodies, a more traditional fedbatch process is typically used. While cells are still growing exponentially, but nutrients are becoming depleted, concentrated feed medium (usually a 10-15 times concentrated basal medium) is added either continuously (as shown) or intermittently to supply additional nutrients, allowing for a further increase in cell concentration and the length of the production phase. In contrast to an intermittent-harvest strategy, fresh medium is added proportionally to cell concentration without any removal of culture broth. To accommodate the addition of medium, a fedbatch culture is started in a volume much lower than the full capacity of the bioreactor (approximately 40% to 50% of the maximum volume). The initial volume should be large enough to allow the impeller to be submerged, but is kept as low as possible to allow for a maximum extension of the cultivation period.

  Fed-batch Culture with Metabolic Shift

  In batch cultures and most fedbatch processes, lactate, ammonium, and other metabolites eventually accumulate in the culture broth over time, inhibiting cell growth. Other factors, such as high osmolarity and accumulation of reactive oxygen species, are also likely to be growth inhibitory, and certainly contribute to the eventual loss of viability and productivity. The effects of lactate and ammonia on cultured cells are complex. Detectable changes in growth, productivity, and metabolism have all been documented. Additionally, metabolite accumulation has been found to affect product quality. In recombinant erythropoietin producing CHO cells, high ammonia concentration has been reported to affect glycoform of the product. By minimizing metabolite accumulation, the duration of a fedbatch culture can be even further extended and higher cell and product concentrations can be achieved. Reduced metabolite accumulation in fedbatch culture is traditionally accomplished by limiting the availability of glucose and glutamine using controlled feeding strategies that maintain glucose at very low levels. After extended exposure to low glucose concentrations, cell metabolism is directed to a more efficient state, characterized by a dramatic reduction in the amount of lactate produced. Such a change in cell metabolism from the normally observed high lactate producing state to a much reduced lactate production state is often referred to as metabolic shift.

  The observation of such changes in metabolism was made more than two decades ago, yet its application in fedbatch culture was not realized until much later. Extending the methodology to controlling both glucose and glutamine at low levels, both lactate ammonium accumulations can be reduced. By applying such a control scheme in fedbatch culture, lactate concentration was reduced by more than three fold, and very high cell concentrations and product titers were achieved in hybridoma cells.

  Continuous Cultures Simple Continuous Stirred Tank Reactor (CSTR)

   Steady state  Grow up the culture in batch mode. Then turn on both in and out flow of medium. Cell and product concentration reach steady state.

   Transient  Same as that for steady state except that cell and product nutrient concentration fluctuate.

  Continuous Culture with Cell Retention (Recycle) – Perfusion Culture

   Transient Same as CSTR, some cells are retained in bioreactor to reach

   high cell concentration. Product throughput is higher per reactor volume, but not the concentration. Typically cell, nutrient and product concentrations fluctuate.

   Steady state  Same as that for transient except that steady state is achieved.

  This rarely happens.

  Continuous Culture with a Metabolic Shift This is the same as simple continuous culture except in the start-up.

  Instead of starting from a batch culture, a fed-batch culture with a metabolic shift is used. After cells reach a high concentration and the metabolic shift is affected, the culture is shifted to a continuous culture. Because no (or low) lactate and ammonia is produced, the concentrations of cells and products are substantially higher than in conventional continuous cultures. In some cases, the cell concentration approaches that of perfusion cultures. However, the medium usage is substantially reduced, and the product concentration is higher.

  Material Balance on Bioreactors Material Balance Equation for Reactor Batch Culture dx _v m

  V = V x _v dt ds V = − Vq x _{s v} dt

  Fed-batch Culture V d x V

  ( d(sv) _v m = F t = x V = − q x V ( ) _{v s v} dt dt dt Fedbatch Culture and Dynamic Nutrient Feeding

  45

of antibiotic production capacity have process engineers played such a key

role in bringing a large array of products to therapeutic use in such a short

time. The increased output required to meet the expanding market was not

accomplished by merely increasing the total culture volume. A large part was

achieved through improving yields by process renovation, as opposed to pro-

cess innovation. Only a decade ago, an antibody titer in the hundreds of

milligrams per liter was the norm. Now, concentrations of a few grams per

liter are common. With the increasing development of new products and the

growing need for large quantities of each new therapeutic, it is prudent to

reassess the technological advances made in the past decade and to pursue

innovative ideas that will ease the task of meeting future demands.

  The final product concentration is primarily affected by the specific pro-

ductivity of cells, the maximum cell concentration, and the duration that

high viability can be sustained. For batch processes, the low level of nutrients

that can be tolerated by cells limits the final cell and product concentration.

Cells are simply unable to attain and sustain high cell concentrations with

the resources available in a typical growth medium. To overcome nutrient

limitation, fedbatch processes have been widely practiced and are currently

the norm for most cell culture processes. In fedbatch cultures, concentrated

medium is added during cultivation to prevent nutrient depletion, prolonging

the growth phase and increasing cell and product concentrations. Continued

addition of medium past the peak of cell concentration also increases the final

titer significantly by allowing cells to be kept viable at high concentrations

and continue to produce product for a longer time.

  Efforts to enhance the performance of fedbatch culture have traditionally

focused on medium development, process control, and manipulation of cell

metabolism by control of the culture environment. With recent advances in

genomic research tools and a more global understanding of cell physiology,

metabolic engineering may emerge as a more prominent strategy to increase

productivity. Even with the promise of creating superior host cells through

cell engineering, pushing the limits of productivity will always require an in-

tensive process engineering effort to accommodate the increased demands of

higher cell and product concentrations. This review will summarize current

practices and articulate the developmental needs of fedbatch culture to meet

these future challenges.

2 Different Forms of Fedbatch Culture

  

Fedbatch processes are widely used in multi-purpose, multi-product facilities

because of their simplicity, scalability, and flexibility. A variety of fedbatch op-

erations, ranging from very simple to highly complex and automated, are seen

in current production facilities. To illustrate the basic operation principles

  46 K.F. Wlaschin · W.-S. Hu

of fedbatch cultures as compared to a typical batch operation, time profiles

of cell, nutrient, and product concentrations for batch (Fig. 1a), intermittent-

harvest fedbatch (Fig. 1b), and traditional fedbatch cultures (Fig. 1c) are

shown.

  In general, fedbatch processes do not deviate significantly from batch cul-

tures. For both intermittent-harvest and traditional fedbatch cultures, cells

are inoculated at a lower viable cell density in a medium that is usually very

similar in composition to a typical batch medium. Cells are allowed to grow

exponentially with essentially no external manipulation until nutrients are

somewhat depleted and cells are approaching the stationary growth phase.

At this point, for an intermittent-harvest fedbatch process (Fig. 1b), a por-

tion of the cells and product are harvested, and the removed culture fluid is

replenished with fresh medium. This process is repeated several times. This

simple strategy is commonplace for the production of viral vaccines produced

by persistent infection, as it allows for an extended production period. It is

also used in roller bottle processes with adherent cells.

  For production of recombinant proteins and antibodies, a more traditional

fedbatch process (shown in Fig. 1c) is typically used. While cells are still

growing exponentially, but nutrients are becoming depleted, concentrated

feed medium (usually a 10–15 times concentrated basal medium) is added

either continuously (as shown) or intermittently to supply additional nutri-

ents, allowing for a further increase in cell concentration and in the length

of the production phase. In contrast to an intermittent-harvest strategy, fresh

medium is added proportionally to cell concentration without any removal of

culture broth. To accommodate the addition of medium, a fedbatch culture is

started in a volume much lower than the full capacity of the bioreactor (ap-

proximately 40% to 50% of the maximum volume). The initial volume should

be large enough for the impeller to be submerged, but is kept as low as pos-

sible to allow for a maximum extension of the cultivation period.

  In batch cultures and most fedbatch processes, lactate, ammonium, and

other metabolites eventually accumulate in the culture broth over time, in-

hibiting cell growth. Other factors, such as high osmolarity and accumulation

of reactive oxygen species, are also likely to be growth inhibitory, and cer-

tainly contribute to the eventual loss of viability and productivity. The effects

of lactate and ammonia on cultured cells are complex. Detectable changes in

growth, productivity, and metabolism have all been documented [1]. Addit-

ionally, metabolite accumulation has been found to affect product quality. In

recombinant erythropoietin producing CHO cells, high ammonia concentra-

tion has been reported to affect the glycoform of the product [2].

  By minimizing metabolite accumulation, the duration of a fedbatch cul-

ture can be even further extended and higher cell and product concentrations

can be achieved. Reduced metabolite accumulation in fedbatch culture is tra-

ditionally accomplished by limiting the availability of glucose and glutamine

using controlled feeding strategies that maintain glucose at very low levels. Fedbatch Culture and Dynamic Nutrient Feeding

  47 Fig. 1

  Representative cell, nutrient, and product concentrations for a typical a batch

culture, b intermittent-harvest fedbatch culture, and c fedbatch culture with dynamic

feeding. As compared to a batch culture, the strategies shown in Figs. b and c extend the

duration and productivity of a culture run by re-supplying depleted nutrients. In fedbatch

culture (c), feed is added continuously to sustain nutrient levels. Much higher cell and

product concentrations are achieved

  48 K.F. Wlaschin · W.-S. Hu

  

After extended exposure to low glucose concentration, cell metabolism is

directed to a more efficient state, characterized by a dramatic reduction in

the amount of lactate produced. Such a change in cell metabolism from the

normally observed high lactate producing state to a much reduced lactate

production state is often referred to as metabolic shift. The observation of

such changes in metabolism was made more than two decades ago [3–7], yet

its application in fedbatch culture was not realized until much later [8]. Ex-

tending the methodology to controlling both glucose and glutamine at low

levels, both lactate and ammonium accumulation can be reduced [7, 9–11].

By applying such a control scheme in fedbatch culture, lactate concentration

was reduced by more than three fold, and very high cell concentrations and

product titers were achieved in hybridoma cells [8].

  Figure 2 compares the time profile of cell growth, glucose concentration

and lactate concentration for two hybridoma fedbatch cultures growing under

different metabolic states. Shown in Fig. 2a is a culture in which the glucose

level was controlled in the range of 1.0–4.0 mM, a relatively low concentra-

tion. In many cultures, glucose concentration is controlled at even higher

levels, in the range of 10 mM. In these ranges of glucose concentration, cells

behave very similarly, having a high lactate production rate. As a result, the

level of lactate accumulated eventually requires the addition of base to main-

tain pH. To supply nutrients to the culture, feed medium was added approxi-

mately proportionally with the base addition rate, since lactate production

is indicative of the metabolic demands of the culture. This feeding strategy

will be discussed in more detail in Sect. 4.2.2. A final cell concentration of

_{6 –1} cells mL was obtained with lactate accumulating to nearly 70 mM

  7.5 × 10

in the final culture volume. In the culture shown in Fig. 2b, the set point of

glucose concentration was at 0.03 mM. Feed medium was added based on

the oxygen uptake rate (OUR), which is estimated on-line. This strategy will

also be discussed further in a later section (4.2.4). The continuous exposure

to very low glucose concentrations allowed cells to shift their metabolism

to a state where little lactate was produced. The final lactate concentration

only accumulated to 40 mM. With the control of glucose concentration at low

levels, the reduced lactate concentration, and the elimination of base add-

_{6 –1} cells mL was ition, a final viable cell concentration of more than 11.5 × 10 achieved.

  Historical data from several batch and fedbatch hybridoma cultures, in-

cluding those shown in Fig. 2, were analyzed to generate the values in Table 1.

Direct comparison of the values between cells in different metabolic states

illustrates that the stoichiometric nutrient consumption and metabolite pro-

duction for cells is notably changed in different metabolic states. Under typi-

cal culture conditions, where nutrients are supplied in excess, more than half

of the carbon in glucose and at least one fourth of the nitrogen in glutamine

consumed is excreted as lactate and ammonium [5, 12]. For hybridoma cells

in a high lactate producing state, this observed stoichiometric ratio is be- Fedbatch Culture and Dynamic Nutrient Feeding

  49 Fig. 2 Time profiles of cell, lactate, and glucose concentration for a hybridoma fedbatch

culture with cells growing with a high-lactate producing metabolism, and b metabolic

shift. Metabolic shift was achieved by control of glucose concentrations at 0.03 mM

tween 1.4–2.2 moles lactate produced per mole glucose consumed. For the

same cells cultured in a metabolically shifted state, a very low ratio of less

than 0.5 moles of lactate produced per mole of glucose consumed is observed.

The ratio of ammonia produced per glutamine consumed is also compared

  50 K.F. Wlaschin · W.-S. Hu

  

Table 1 Characteristic Stoichiometric Ratios of Key Nutrients for Cells Growing in Differ-

ent Metabolic States Stoichiometric ratio Without Metabolic shift Lactate (mmole/mmole) metabolic shift consuming cells lactate/glucose 1.4 – 2.2 0.05 – 0.5 0.4 – 1.0

• – ammonia/glutamine 0.5 – 1.3 0.1 – 0.3
• – alanine/glutamine 0.2 – 1.3 0.01 – 1.3 1.0 1.0 – 2.0
• – oxygen/glucose

  

in Table 1, showing a dramatic reduction from 0.5–1.3 moles ammonium

per mole of glutamine to 0.1–0.3 mole per mole under metabolically shifted

conditions. In later stages of fedbatch cultures, lactate consumption, as op-

posed to production, is occasionally observed, although this phenomenon

is not well documented in published literature. In such cases, an approx-

imate ratio of lactate to glucose consumption is between ∼ 0.4–1.0 moles

of lactate consumed per mole of glucose consumed. While this observation

seemingly contradicts the role of lactate as an inhibitory molecule, it illus-

trates the flexibility of mammalian cells to adapt their behavior for survival

under a wide range of conditions. With this repertoire of available cell behav-

ior, fedbatch culture strategies that provide conditions that reduce metabolite

accumulation is a field of fedbatch culture technology still warranting further

development.

3 Designing Feed Medium for Fedbatch Cultures

  

The design of feed medium is critical for the implementation of a success-

ful fedbatch process. A well-designed feed medium should ensure cell growth

and product formation are not limited by depletion of any medium compon-

ent or inhibited by excessive nutrient concentration or metabolite accumula-

tion. To achieve this, a good estimate of the rates of consumption of medium

components is required. For most processes, a feed medium that is 10 to 15

times the nutrient concentration of basal medium is used. With this simple

design, the consumed nutrients are replenished, and the growth and produc-

tion phases are prolonged; however, many components will likely be supplied

in excess, while others will be in limited supply [13].

  The nutritional requirements for mammalian cells are very complex. Most

media contains glucose, vitamins, and virtually all amino acids. Among the

amino acids included, 13 are deemed “essential” for cultured cells, as most

cell lines cease to grow in their absence [14, 15]. This requirement for cul-

tured cells is higher than the 11–12 essential amino acids required for survival