lecture 4 - physical bioprocesses 2

7/29/2019 Lecture 4 - Physical Bioprocesses 2

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Physical Processes in Bioreactors 1

Master of Sciencein Chemical Engineering

Course in

University of San CarlosDepartment of Chemical Engineering

Nasipit, Talamban, Cebu CityTelefax: +63323446783

Email: [email protected]

Bioprocess Technology

Physical Processes in Bioreactors


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Contents of the Lecture

Fluid flow and mixing

Heat transfer

Mass transfer


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Heat Transfer in Bioreactors

In situ sterilization of liquid medium Temperature control during reactor operation

Heat-transfer configuration for bioreactors

a) Jacketed vessel

b) External coil

c) Internal helical coild) Internal baffle-type coil

e) External heat exchanger


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Heat-transfer configurations for

bioreactors


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Temperature changes for control of fermentation

temperature using cooling water

Te m

p e r a t

u r e

Distance from cold-fluid inlet

TF

Tci

Tco

Fermenter temperature


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Heat Transfer Mechanisms

RECALL, Conduction

Convection Radiation


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Heat Transfer Between Fluids

ThThw

Tcw

Tc

Hot fluidCold fluid

(1) (2) (3)

(1) ( )H h h hwh A T T =

(2)

( )

( )hw cw

H

T T

L

=

(3) ( )H c cw ch A T T =

Resistances: 1h

h

R

h A

=

1c

c

Rh A

=

w

LR

=

Where:

h = individual heat transfer coefficient.


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Overall Heat Transfer Coefficient

( )H h cUA T T = (a)

1T h w cR R R R

UA= = + +

But,

1 1 1

h c

LUA h A A h A

= + +


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Fouling Factors

1 1 1 1 1fh h c fc

LU h h h h

= + + + +

Fouling factor of the hot-fluid sideFouling factor of the cold-fluid side

Design equations for heat-transfer systems

Equation (a) and energy balance equations


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Estimation of Heat Transfer

Coefficients

= wb

L

D

GrfNu

,,Pr,Re,


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Heat Transfer in Aerobic

Fermentations

rate of oxygen consumption by cells

Maximum cell concentration supported by heat transfer system

rxnH


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Heat Transfer in Aerobic

Fermentations

9Heat transfer is severe when: Fluid is viscous

Turbulence & high U are difficult toachieve


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Mass Transfer in Bioreactors

9Conduction (role in bioprocessing) Scale of mixing

Solid-phase reaction Interphase transfer

9Convection


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Diffusion TheoryCA1

CA2

JA

x

dx

dC

Da

N

JA

AB

A

A ==


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Mass Transfer Across A Phase Boundary

=

propertiesfluid

geometry,ics,hydrodynamfk

Rate of transfer

AA

CkaN =

Resistance to mass transfer

kaRm1

=


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Mass Transfer Across A Phase

Boundary

Liquid-solid Liquid-liquid

Gas-liquid


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Liquid-solid Mass Transfer


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Gas-liquid Mass Transfer


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Oxygen Transfer


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Steps for Oxygen Transfer from gas bubble to cell

i. Transfer from the interior of the bubble to the gas-liquidinterface

ii. Movement across the gas-liquid interfaceiii. Diffusion through the relatively stagnant liquid film

surrounding the bubble

iv. Transport through the bulk liquid

v. Diffusion through the relatively stagnant liquid filmsurrounding the cells

vi. Movement across the liquid-cell interface

vii. If the cells are in a floc, clump or solid particle,diffusion through the solid to the individual cell

viii. Transport through the cytoplasm to the site of reaction


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Analysis of G-L Mass Transfer in Bioreactors

i. Transfer through the bulk gas phase in the bubble is relatively fastii. The gas-liquid interface itself contributes negligible resistance

iii. The liquid film around the bubbles is a major resistance to oxygen transfer

iv. In a well-mixed fermenter, concentration gradients in the bulk liquid are

minimized & mass-transfer resistance in this region is small. However, rapidmixing can be difficult to achieve in viscous fermentation broths; if this is the

case, oxygen-transfer resistance in the bulk liquid may be important.

v. Because single cells are much smaller than gas bubbles, the liquid film

surrounding each cell is much thinner than that around the bubbles and its

effect on mass transfer can generally be neglected. On the other hand, if thecells form large clumps, liquid-film resistance can be significant.

vi. Resistance at the cell-liquid interface is generally neglected.

vii. When the cells are in clumps, intraparticle resistance is likely to be significant

as oxygen has to diffuse through the solid pellet to reach the interior cells.

The magnitude of this resistance depends on the size of the clumps.

viii. Intracellular oxygen-transfer resistance is negligible because of the small

distances involved.


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Oxygen demand during cultivation of yeast cells

In the culture medium the O2 solubility is

CO2* = 1.1 mmol/l = 35.2 mg / 1 (ppm),

(a)

Which in case of air saturation reduces to

CO2* = 35.2 x 0.209 = 7.36 ppm

(b)

The typical O2 demand of the yeast fermentation is in the following range

Q O2 = 0.3 g O2 / (g Dm h) (c)

Hence, for biomass concentration of 20 g Dm/l the oxygen demand is

Q O2 = 0.3 x 20 = 6 g O2 / h

(d)This implies that the number of complete saturations of the fermentation broth amounts to

n-Sat = 6 / 0.00736 = 816 h-1

(e)

This is not an easy task, particularly, if higher biomass concentrations are desired which are usual inpractice (for instance 60 g DM/l).

Critical oxygen concentration


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Critical oxygen concentration

To avoid limitations of cell growth and product formation by lack of oxygen it is necessary that a certain

oxygen concentration is maintained in the culture media.

At next, an estimation is done to find out the conditions under which the biochemical reactions involved in

cell growth are the limiting steps. Usually, the highest oxygen consumption rates are observed during cell

growth. The largest mass transfer rates for oxygen (mol/cm3.s) are obtained for CL = 0, hence

max = kL a CL (1

On the other hand, the maximum oxygen consumption during cell growth is given by

rmax

= max X . 1/YO2 (2)

with the yield coefficient

YO2 = generated biomass / consumed moled O2 (3)

It is understood that the main resistance for oxygen utilization is predominantly limited by microbialmetabolism if

max >> rmax (4)

or

kLa CL* >> 1/YO2 max X , (5)

respectively.


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In this case there is no mass transfer resistance for oxygen and the entire growth process is controlled

by the biochemical reactions in the cell machinery. However, if CL < CL, crit

( = Ko2 ), oxygen can be the limiting substrate and hence the oxygen input into the culture media is thelimiting step of the process (Fig. 9).

With the assumption CL* >> CL it follows

X/'ak*CY1

X/'akKY

maxLLO

maxOO

2

L22

It can be anticipated as a guideline that no mass transfer limitation for oxygen exists at the gas/liquid

interface if eq. (38) appliesC

L> 3

2OK ( = CL, crit )

From the biochemical viewpoint this implies that always sufficient oxygen is available to except the

electron pairs which pass through the respiratory chain. To estimate the critical oxygen concentration

which leads to growth limitation by lack of oxygen a kinetic law of Monod type is assumed to apply

for oxygen, i.e.,

r = max

x2

2OLO

L

Y

1

CK

C

+

Then the stationary oxygen balance is given by

OTR = OUR

Oxygen transfer rate oxygen uptake rate

kL

a (CL* - C

L) =

maxx

22

OLO

L

Y

1

CK

C

+

(6)

(7)

CL = CL* (8)

(9)


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Typical values of the critical O2

concentration

(at sufficient supply with other substrates)

CL* (at air saturation) = 0.23 mmol/l

Microorganisms T, C CL,

Crit mmol/l

Azetobacter vinlandii 30 0.03

E. coli 37.8 0.008

15 0.003Pseudomonas denitroficans 30 0.009

Yeast 34.8 0.004620 0.0037

P. chrysogenum 24 0.022

30 0.009It can be assumed

as a rule of thumb that0.003 < C

L, crit < 0.05 mmol/l

which corresponds to an air saturation of 1 to 25%

Fig.9


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Scale-up

The performance of a bioreactor depends on process specific data (stoichiometric coefficients,

thermodynamic and kinetic data) and transport parameters (hydrodynamics, heat and mass

transfer properties).

Specific data are scale independent and the transport parameters may be drastically dependenton the reactor size. For instance, the oxygen solubility (CL*) is a thermodynamic quantity

which depends only on the medium composition and the temperature while the actual oxygen

concentration in the culture medium is a complex function of the biochemical kinetics and the

scale dependent transport parameters (kLa, , etc.).

Phenotypical appearance of an organism is determined by both its genotype and the living

conditions in the culture.

Kossen has proposed to select production strains in laboratory scale at conditions which prevailin large reactors. Instead of looking for a baseball player in a soccer field, one should do scale

down of reactors for proper screening of microbes.


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The different methods for scale-up can be summarized as follows:

(1) Fundamental methods (modeling of the process on the basis of the balances for mass, heat

and momentum which requires detailed knowledge of all transport effects and their

interaction)

(2) Semifundamental methods (solution of simplified balance equations)

(3) Dimensional analysis (is only of limited value as not always applicable)

(4) Rules of thumbs, know how(5) Trial and error

In todays fermentation industry only methods (4) and (5) are established. The often used rules

of thumbs imply: Do the scale-up in such a way that the culture conditions (environment)

and some typical engineering parameters (P/V, , kLa, Re, ) are kept constant.

Scale-up criteria used in industry

% of industry Criterium used

30% P/V constant30% kLa constant

20% utip constant*

20% Po2 constant

*utip = stirrer tip velocity


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Examples for a successful scale-up on the basis of different criteria are given in Fig. 10.

The intrinsic know how is the choice of the correct criteria for the specific process.

Different scale-up criteria and their consequencesScale-up from a 10 litre reactor to 10 m3 which is a linear scale factor of 10.

1 0.22 2.15 21.5

102 1 10 102

0.1 0.1 1 10

10-4 10-2 0.1 1

103

105

102

0.1

P/V

N(-1)

utipRe

Value in 10 m3 scale

P/V N(-1) ND Re

PCriterium:

Constancy of


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Some criteria cannot be fulfilled in practice. An example is the requirement for equal mixing times in

small and large scale. For many stirrer systems we have the empirical relation

300D

p5

3

=

Considering this, the equality of mixing times ( = idem) implies that

l2S2)

D

V/P()

D

V/P( =

2

S

l

)D

D

For a linear scale factor of 10 (0.01 to 10) and (P/V)S

= 2 kW/m3 it follows

(P/V)l= 200 kW/m3

Such a high energy input in a bioreactor is nearly impossible and would be extremely costly.

(P/V)l = (P/V)S (

(10)

(11)

(12)

lecture 4 - physical bioprocesses 2

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