TECHNICAL PAPER Stirred Bioreactor Engineering for Production Scale, Low Viscosity Aerobic Fermentations: Part 1 By: Dr. Alvin Nienow Senior Technical Consultant The Merrick Consultancy Merrick & Company 2450 S. Peoria Street Aurora, CO 80014-5475 Tel: 303-751-0741 Fax: 303-751-2581 www.merrick.com January 2012
TABLE OF CONTENTS Section Introduction................................................... 1 Engineering Issues in Stirred Bioreactors, Part 1.......................... 2 Mass Transfer of Oxygen into the Broth and Carbon Dioxide out....................... 3 Hold-up and Foam Formation.............................................. 4 Heat Transfer........................................................ 4 Summary..................................................... 5 References.................................................... 6 Nomenclature.................................................. 7 Tables....................................................... 8 Figures...................................................... 9 REMAINDER TO COME Page Alvin W. Nienow Merrick Consultancy i
Introduction There are many examples of important bioprocesses that fall into the category of low viscosity aerobic fermentations. They include the use of genetically-modified bacteria such as Escherichia coli to give bulk chemicals such as 1,3-propanediol; and E. coli and Pichia pastoris for the production of proteins for medical purposes. Indeed, though biofuels such as bioethanol, biobutanol and biogas are made under anaerobic conditions, biodiesel can also be produced aerobically from genetically-modified E. coli. Other examples include Saccharomyces cerevisiae for the manufacture of Baker s yeast and Corynebacter glutamicum for valine and lysine production as animal feed additives. The bioreactor (fermenter) of choice for such processes uses mechanical agitation and a typical traditional stirring configuration based on Rushton turbines is shown in Fig 1a in diagrammatic form. Such fermenters have been in use for at least 50 years and are commonly up to 250 m 3 in scale or even larger for the manufacture of many commercially important products. Fig 1b shows a photo with some people inspecting the internals in a smaller scale industrial bioreactor. In this case, the impellers are of the wide-blade, high solidity ratio type which is one of the type more commonly used today as a result of much research which has shown that changing the agitator from the Rushton turbine to newer types can lead to significant improvements in the fermentation process. In these series of articles, the mixing issues which have to be considered when designing or operating such fermenters will be discussed. In particular, the reasons why modern impellers are replacing the traditional Rushton turbine by retrofitting into extant industrial equipment or have become the impellers of choice when new plant is built will be explained. These impeller developments have been largely ignored in chemical and biochemical engineering textbooks and in general even in the refereed journals except for a few specialist exceptions. It is also not the intention to give detailed design details or methods of calculation. These aspects are adequately covered in standard textbooks, especially that by Van t Riet and Tramper, Basic Bioreactor Design 1. The essence of such decisions comes from recognising that the issues can all be considered as bioprocess scale-up where information obtained generally on a much smaller scale in the laboratory or pilot plant is used to establish the desired commercial scale operating conditions. In other words, the industrial scale bioreactor has to provide a suitable environment for the organism to grow and produce, based on the work conducted on this smaller scale. In general, it is not possible to mimic on the industrial scale exactly the conditions found in the smaller scale, so inevitably scale-up is a compromise. Essentially, however, this compromise is best understood if it is recognised that each cell is itself a mini-factory converting nutrients, generally carbon based, into the desired product. Thus the total production rate depends on production rate per cell multiplied by the number of cells in the fermenter times the size of the fermenter. If the overriding importance of providing the correct environment for the cell is recognised, scale-up, though indeed a compromise, is based on that concept rather than arbitrary scale-up rules related to fluid dynamics, which essentially ignore the well-being of the cell. That is the approach recommended in this series of articles. The equipment used to obtain this information on the desired environment for the cell is indicated in Fig 2. At the two smallest laboratory scales, the shaken microwell or shake flask (Fig 2 a and b) are used. The biological parameters that may be determined in them are shown in Table 1. To find the optimum, many experiments must be undertaken to establish the correct value of each of the many parameter; and this type of equipment is ideal for such a requirement. It is also useful to establish the sensitivity to variations from that optimum as such variations are bound to occur on the commercial scale. These variations arise because the environment in the 250 m 3 stirred reactor is clearly not going to be as spatially homogeneous as the small volumes found in the shaken microwell and shake flask. One of the big advances that has been made in recent years in developing bioprocesses has been the ability to both measure and control many of these parameters in each reactor even at these small scales. The papers of Buchs and co-workers, starting with one from 2011, 2 give an excellent indication of the developments in this area of allowing quantitative data for engineering purposes to be obtained at these small scales. As also shown in Table 1, some other biologically-specific parameters must be determined at the larger stirred bench scale (Table 1b) as shown in Fig 2c in order to be useful. Fed-batch bioreactors are ones in which additional growth medium is added over time, thereby increasing the cell concentration that can be achieved (as well as the volume of medium). Thus, though the amount of oxygen that each cell requires to function properly can be determined at the very small scale, to get a satisfactory indication of how that oxygen will be supplied Alvin W. Nienow Merrick Consultancy 1
in the commercial plant especially under fed-batch conditions really requires experiments to be done in similar equipment. Essentially, that is linked to the need to use stirring to provide the energy for oxygen transfer with air sparged from the base (specified by the mass transfer coefficient, k L a, as set out in more detail later) rather than by shaking and from headspace aeration respectively. The potential for the energy input from the impeller to damage the organism, often referred to as shear damage can also be assessed in such relatively small scale stirred bioreactors. This possibility comes about because though the bench scale bioreactor is small compared to the commercial one, even at this scale, the turbulence in the flow (which is discussed further below) that impacts on the cell is similar at the small scale and the large. So the size of the cell to the scale of the turbulent flow is also similar. Having the information on the bioprocess of interest at the scales indicated in Figs 1a to 1c, good engineering must then be used to determine suitable conditions at the production scale. These conditions should match sufficiently closely those shown to be optimum at the bench scale with respect to cell and product concentration. Thus, the desired temperature, ph, dissolved oxygen concentration, etc, must be achieved by the agitation and aeration system in the bioreactor. The agitator provides the energy which gives rise to the turbulent liquid motion required for these conditions to be achieved. In particular, the dispersion of the air must be done so that the oxygen transfer rate, OTR, is able to meet the oxygen required by the cells to function properly. The turbulent flow also homogenises the contents of the fermenter so as to maintain the various operating parameters within the range that ensures the cells produce a similar bioprocess performance at the commercial scale to that obtained at the small. To give confidence that the commercial plant is going to achieve similar results to the smaller scales, operation at the pilot plant scale shown in Fig 2d is also often undertaken. This overall approach is bioprocess scale-up with the cell and its local environment as the key. The aim of much current research is to establish ways of going from as small a scale as possible to the commercial scale whilst minimising work at intermediate scales, in particular to eliminating the pilot scale stage. This approach is greatly enhanced if the smaller scales, especially in the stirred bench scale bioreactor, are conducted in such a way that important parameters such as specific power (W/kg) are similar to those that will be used on the large scale. Choosing sensible parameters in this way for the bench scale is often called scale-down. A good scale down protocol greatly eases scale up The remainder of these articles aim to help understand the interaction between the fluid motion generated by different agitators under aerated conditions with variations in speed and power input; and how they change with scale. This understanding is important for the design of new equipment; and retrofitting and solving operational issues on that already extant. To aid this understanding it is useful to subdivide this overall task into a sub-set of smaller issues. These issues will now be discussed. Engineering Issues in Stirred Bioreactors Many of the engineering issues are generic and apply to all aerobic bioprocesses (and indeed gassed chemical reactors in general). They can be considered as physical parameters ; and those most relevant to bacterial fermentations are listed in Table 2. Table 2a sets out the quantitative parameters required for design whilst Table 2b lists those parameters which aid understanding and have helped improve large scale operation and design. The parameters in Table 1 on the other hand are specific to the organism being grown and will usually be different for each case. Achieving satisfactorily the biological parameters required by the cell in Table 1 by judicious design and appropriate selection of the parameters in Table 2 is essentially the task of the engineer, especially with respect to scale-up. The final process engineering specification of the fermenter will be the size (diameter, T (m), height, H (m)), the impeller type(s), number and size, D (m), the power needed to be imparted to the broth, P (W) and the aeration rate, Q G (m 3 s -1 ). The way these values and the parameters set out in Table 2a are determined will now be discussed. For the microbial fermentations being considered here, the viscosity of the growth medium with the cells in it (the fermentation broth) essentially does not go much higher than that of water. With such low viscosities, the flow in the fermenter is turbulent at the 5 L bench, i.e. Reynolds number, Re =þnd 2 /µ > ~ 10 4 where þ is the broth density (kg m -3 ), µ, its viscosity (Pa s), D,, the impeller diameter (m) and N, its speed (rev s -1 ). In practice, Re increases with increasing scale. However, as the flow is turbulent, the actual value of the Reynolds number does not matter and turbulent flow theories can be used to analyse the fluid mechanics in the bioreactors across the scales. The topics listed in Table 2 will now be considered for such flows. Alvin W. Nienow Merrick Consultancy 2
Mass Transfer of Oxygen into the Broth and Carbon Dioxide out The transfer of oxygen from air into a fermentation broth has been used since the 1940s when depth fermentations were first established. It is one of the most important aspects of fermenter operation because oxygen is only sparingly soluble in water and therefore in the medium, which largely consists of water. If the supply of oxygen into the broth ceases, its concentration in the broth would generally fall below the desired value in less than a minute. Thus, the overall oxygen demand of the cells throughout the batch or fed-batch fermentation must continually be met by the oxygen transfer rate, OTR (mol O 2 m -3 s -1 ); and the demand increases as long as the number of cells is increasing. Thus, a maximum oxygen transfer rate must be achievable and this depends on the mass transfer coefficient, k L a (s -1, though units of min -1 are also often used), and the driving force for mass transfer, C L (mol O 2 m -3 ). Thus, OTR = k L a. C L 1 For oxygen transfer, the driving force, C L, conceptually is the difference between the oxygen concentration in the liquid film around the air bubbles (the concentration in equilibrium with the partial pressure of oxygen in the bubble) and that in the broth. The latter concentration must always be held above a critical do 2 value as determined at the well-mixed bench scale. This concentration is usually expressed as a % of saturation with respect to air as measured by a dissolved oxygen probe (% do 2 ). Though the critical value is often less than 10% do 2 (and sometimes close to zero), the set value is often as high as 40% do 2. This higher value is to ensure that do 2 never falls below the critical value throughout the fermenter since spatial homogeneity is difficult to achieve at the commercial scale as discussed qualitatively above and quantitatively later. In practice, more oxygen transfer can be achieved by increasing the driving force by raising the O 2 partial pressure in the incoming gas stream either by imparting a back-pressure on the bioreactor or by adding extra oxygen to the sparged air, preferably as a separate flow. Roughly, for every mole of O 2 taken up by a cell, 1 mole of carbon dioxide, CO 2, is produced. This ratio of moles CO 2 produced to O 2 consumed is called the respiratory quotient, RQ, and as suggested here it is generally ~ 1. Because CO 2 is very soluble in the broth, especially compared to oxygen, it dissolves. The value of k L a is similar for both O 2 transfer from air to the broth and CO 2 from it; but the high solubility, CO 2 makes it much more difficult to strip out. However, stripping of dissolved CO 2 is very important because above a certain value of dissolved CO 2 (pco 2 > ~ 100 mbar), a reduction in fermentation rate or productivity is generally observed. In addition to the driving force, the other parameter that the engineer can manipulate to control the rate of mass transfer is the k L a. In low viscosity systems, k L a is only dependent on two parameters. These are, firstly, the power input, P (W), into the fermenter, mostly from the impeller, and the air flow rate, Q G (m 3 s -1 ). It has been shown that if the power input from the impeller when air is being sparged, P g (W) is normalised in terms of the kg of broth, P g /M (W kg -1 ) in the fermenter, k L a can be correlated with this parameter at different scales. The specific power input into the fermenter from stirring is numerically equivalent to the mean specific energy dissipation rate, (ε T (W/kg); and because this is a fundamental parameter in the modern understanding of turbulence, it will be used here in the rest of these articles. Thus, P g /M / (ε T. Generally, for economic and biological reasons for these types of fermentation, the specific power (mean specific energy dissipation rate) from the impeller is about 1 to 5 W kg -1. The airflow rate impacts in the same way across the scales if the superficial air velocity, v S (m s -1 ) through the fermenter is used, where v S = Q G /A where A T (m 2 ) is the cross-sectional area of the fermenter and A T = πt 2 /4 where T (m) is the diameter of the fermenter. However, the amount of broth and the number of cells in the fermenter increases in proportion with its volume. As a result so does the volume of oxygen required and the amount of carbon dioxide produced. Therefore in order to satisfy the mass balance for O 2 transfer in and CO 2 out, the volumetric flow rate of air needs to be kept essentially constant. If this flow rate is expressed on a per minute basis, a typical practical value is about 1 vvm (where vvm is the volumetric flow rate of air at standard conditions in m 3 min -1 per m 3 broth in the bioreactor). Thus, v S = (vvm/60)(volume of broth, m 3 )/(X-sectional area of the bioreactor, m 2 ) 2 The combination of (ε T and v S selected must together be sufficient to produce the necessary k L a where k L a = A(ε T ) a g (v S ) b 3 k L a is difficult to determine experimentally especially on the large scale but within the accuracy achievable, this equation is found to be independent of the number of impellers and their type and also of scale 3 ; and a and b Alvin W. Nienow Merrick Consultancy 3
are usually about 0.5 ± 0.1 for low viscosity broths. On the other hand, the numerical value of A (which is NOT dimensionless) is extremely sensitive to composition. Thus, a typical value of k L a would be about 0.1 to 0.2 s -1 in water with the addition of antifoam lowering it by a factor of up to 2; and salts increasing it up to x4 for the same values of (ε T and v S. It is important to point out at this stage that though Equ 3 does not depend on impeller type, the value of (ε T and v S that can be efficiently utilised does. Thus, impeller choice is very important if the mass transfer requirements are to be met and this aspect will be discussed in the next article. As pointed out earlier, the value of k L a is similar for both O 2 transfer in and CO 2 transfer out. Thus, provided scale-up is undertaken at constant vvm (or close to it), the driving force for transfer of O 2 and of CO 2 will remain essentially the same across the scales. As a result, the total volumetric gas flow rate into the bioreactor should also be able to strip out the CO 2 to give the same partial pressure of CO 2 in the exit gas and pco 2 in the broth on scale-up as on the small scale, thus preventing potential problems with high values of this parameter. In addition, as the air volumetric flow rate scales with fermenter volume, at constant vvm, v S increases with the linear scale-up ratio (approximately with T commercial scale /T bench scale ), thereby enhancing k L a if constant (ε T scale-up is also used. Hold up and Foam Formation There is a down side to the higher superficial velocity on scale-up. These higher superficial gas velocities increase hold-up (hold-up is essentially the proportion of the fermenter taken up by gas bubbles of different O 2 concentration). Higher hold-up means loss of fermenter capacity as cells are only producing in the broth and not in the gas phase. Even more problematic is that the high v S increases the tendency to form a stable foam, which further lowers productivity and if not controlled, may lead to broth being driven out of the bioreactor into the exit pipe, in extreme cases causing shut down. The usual way of handling such problems is to use one of a variety of anti-foams 4. Their use has two disadvantages. They are expensive and, as mentioned above, they lower k L a. It has been shown that the certain modern impellers of the type discussed later reduce the foaming tendency compared to Rushton turbines 5. Though not so well documented, experience at the industrial scale suggests that retrofitting to these other impellers has increased k L a by about 20 to 30% at the same (ε T and v S because less antifoam has been used. If retrofitting is carried out, monitoring the use of antifoam is a useful way of establishing the change in running costs and also perhaps explaining the higher k L a that may well be found in practice. Heat Transfer Accurate temperature control is very important in fermentation processes as cells are very sensitive to that parameter which should normally be held between about 35 to 40 C. The oxygen uptake rate largely determines the metabolic heat release Q H (W kg -1 ) so that Q H 4.6 x 10 2 OUR 4 This cooling load has to be removed by heat transfer at an equivalent rate given by Q H.M = U A H u 5 where U is the overall heat transfer coefficient (typically about 2000 to 3000 W m -2 C -1 ), u ( C) is the difference between the temperature of the cooling water and the broth temperature (about 35 to 40 C) and A H (m 2 ) is the heat transfer area available. U is hardly affected by the agitation conditions though larger D/T impellers associated with lower power number, Po, impellers (as discussed later) maximise the inside heat transfer coefficient for a given (ε T. Overall, at the commercial scale, heat transfer is often a problem as the cooling load scales with the volume of the reactor, i.e., approximately reactor diameter T 3 whilst cooling surface area scales with T 2. Thus, if only a cooling jacket as shown in Fig 3a is used on the larger scale, the area/volume goes down compared to the bench scale so that cooling coils are often required and sometimes cooling baffles (Fig 3b). Inability to meet the cooling requirements at the large scale is a very serious problem because it is extremely expensive to resolve and cannot be achieved by increasing the agitation intensity as the overall heat transfer coefficient, U is insensitive to it. Thus, if it is necessary to increase Q H because it is insufficient to meet the cooling requirements of the fermentation, either A H or u must be increased. The former requires the whole fermenter construction to be modified and the latter needs refrigeration of the cooling water. It is clearly better to design with a high safety margin with respect to the area in the first place. In a particular industrial example with which I was involved, the rate of oxygen transfer was increased to meet the oxygen demand of the organism by introducing a slow oxygen feed through a separate sparger (which incidentally is the best way of using oxygen to give an enhanced driving force). That approach was very successful as a way of enhancing oxygen transfer but the heat release was such Alvin W. Nienow Merrick Consultancy 4
that the temperature could no longer be controlled. As a result, the nutrient feed to the fed-batch fermentation had to be slowed down below the maximum rate now achievable with the higher rate of oxygen transfer. Summary In the Introduction, the need to consider the cell as the productive source in a bioreactor is introduced. In Part 1, those engineering parameters which have to be satisfied if cells are to grow satisfactorily at the large scale are discussed. In particular, they are the provision of oxygen in the broth at an appropriate concentration; and the need to strip out to a low concentration, the carbon dioxide that is produced at a rate governed by the rate at which oxygen is consumed by the growing cells. The critical parameters in these two processes are the specific power input from the stirrers and the airflow rate, of which increases in both increase the mass transfer coefficient and of the latter, the driving force for mass transfer of O 2 and CO 2. The air flow rate also impacts on the amount of gas held up in the bioreactor and the tendency for foam to form, both of which can reduce its productive capacity. Finally, the heat evolved at a rate determined by the rate the organism takes up oxygen needs to be removed in order for the bioreactor to be cooled in order to operate at the desired temperature is discussed. The need to be rather conservative with the area available for cooling at the large scale is emphasized. In Part 2, the modern impellers that have been introduced into industry in recent years will be discussed. In particular, their impact on two aspects will be emphasized. The first topic will show how improved mass transfer can be achieved by the correct choice of impellers, thereby obtaining a higher volumetric productivity from a bioreactor. The second topic will show that there is an inevitable temporal and spatial increase in the range of temperatures and concentrations of nutrients, dissolved oxygen, ph, etc., experienced by the cells on the larger scale even if the mean value is closely controlled to coincide with that determined at the small scale. However, by a suitable feed strategy and choice of impellers, it will be seen that these variations can be greatly reduced, again giving an improved performance. Alvin W. Nienow Merrick Consultancy 5
References 1. Van t Riet, K and Tramper, J, Basic Bioreactor Design, Marcel Dekker, Inc., New York, USA, 1991. 2. Huber, R, Roth, S, Rahmen, N and Büchs, J. Utilizing high-throughput experimentation to enhance specific productivity of an E.coli T7 expression system by phosphate limitation. BMC Biotechnology, 11, (2011), 22-33. 3. Nienow, AW, Scale-Up, Stirred Tank Reactors. In: Encyclopedia of Industrial Biotechnology, John Wiley & Sons, Inc., Hoboken, NJ, USA, DOI: 10.1002/9780470054581.eib535: Vol. 7 (2010) 4328-4341. 4. Nienow, AW, Aeration-Biotechnology, In: Kirk Othmer Encyclopedia of Chemical Technology, 5th Edition, Wiley, New York, USA, 2003. 5. Denkov, ND, Mechanisms of Foam Destruction by Oil-Based Antifoams, Langmuir, 20, (2004), 9463 9505. 6. Boon, LA, Hoeks, FWJMM, van der Lans, RGJM, Bujalski, W, Wolff, MO and Nienow, AW, Comparing a Range of Impellers for Stirring as Foam Disruption (SAFD), Biochem. Eng. J., 10, (2002), 183-195. Alvin W. Nienow Merrick Consultancy 6
Nomenclature A H, heat transfer area A T, cross-sectional area of the fermenter AR, aspect ratio, H/T B, baffle width C, bottom impeller clearance above vessel base D, agitator diameter H, bioreactor fill level k L a, specific mass transfer coefficient M, mass of broth in the fermenter N, agitator speed OTR, oxygen transfer rate from the gas phase OUR, oxygen uptake rate by the cells P, power Q G, volumetric gas flow rate Q H, metabolic heat evolution rate Re, Reynolds number (= þnd 2 /μ) RQ, respiratory quotient T, bioreactor diameter U, overall heat transfer coefficient v S, superficial gas velocity vvm, specific volumetric air flow rate Greek Letters a, b, exponents C L, driving force C, spacing between impellers with dual or more impellers u, temperature driving force ε T, local specific energy dissipation rate ε T, mean specific energy dissipation rate μ, viscosity v, kinematic viscosity þ, liquid density u m, mixing time Subscripts g when air is sparged Alvin W. Nienow Merrick Consultancy 7
Tables Table 1 Bioprocess Specific Data (Nienow, 2010) a) That Obtainable in Shaken Microtiter Plates or Shake Flasks 1 Media design 2 Metabolite concentrations including product inhibition* 3 Feeding algorithm for fed batch* 4 Choice of ph control agents and sensitivity to ph* 5 Temperature sensitivity* 6 Growth and production patterns * 7 Oxygen-demand, CO 2 production and RQ profile * 8 Heat-release rate (probably estimated well enough from OUR) * 9 Substrate utilization efficiencies * 10 Cell and product concentrations * 11 do 2, pco 2 and osmolality tolerance* * The main problem is the measurement and control of ph, do 2, pco 2 and cell mass at these scales. b) That Obtainable from Batch/Fed-Batch Bench Scale Stirred Bioreactors 12 k L a profile ( including impact of antifoam) 13 Foaming/hold-up characteristics 14 Sensitivity to fluid dynamic generated stresses from agitation or bubbling aeration 15 Stresses associated with spatial broth inhomogeneity on scale-up Table 2 Generic Physical Parameters Required for Design/Scale-up (Nienow, 2010) a) Parameters most essential for scale-up/design 1 Adequate rate of mass transfer (O 2 in, CO 2 out)* 2 Bubble hold-up* 3 Satisfactory heat transfer for temperature control* 4 Impeller power number, Po, and unaerated motor power draw, P = PoþN 3 D 5 5 Mean specific energy dissipation rate, ε T = P/M W kg -1 6 Reduced power draw on aeration, P g ((ε T = P g / M) 7 Good air dispersion 8 Effective bulk fluid mixing * These values will be specific to each bioprocess b) Parameters which aid understanding 9 Flow close to the agitator single - and two phase air-liquid 10 Spatial variation in local specific energy dissipation rates, ε T W kg -1 11 Gas phase mixing Alvin W. Nienow Merrick Consultancy 8
Figures Figure 1a. Schematic representation of multiple Rushton impellers in fermenters (D/T =1/3; B/T =1/10; C/T = 1/4; H/T = ~ 3; 4 baffles) Alvin W. Nienow Merrick Consultancy 9
Fig 1b A photo taken during the installation of wide blade bydrofoil impellers in a fermenter. Alvin W. Nienow Merrick Consultancy 10
a) Shaken micro-titer plate b) Multiple shake flasks Alvin W. Nienow Merrick Consultancy 11
c) Bench scale fermenter d) Pilot scale fermenter Fig 2 Examples of the range of scales used to establish the required commercial operating conditions. Alvin W. Nienow Merrick Consultancy 12
Wide blade hydrofoil Cooling Jacket Hollow blade impeller a) Cooling coil Baffles which may also be used for cooling b) Fig 3 Cooling surface area provided by a) a jacket; b) coils Alvin W. Nienow Merrick Consultancy 13