Bioreactor Size

000 to 10,000 plantlets from plant tissue, which can then be transplanted directly into soil. 

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The above correlation is valid for a bioreactor size of less than 3000 litres and a gassed power per unit volume of 0.5-10 kW. For non-coalescing (non-sticky) air-electrolyte dispersion, the exponent of the gassed power per unit volume in the correlation of mass transfer coefficient changes slightly. The empirical correlation with defined coefficients may come from the experimental data with a well-defined bioreactor with a working volume of less than 5000 litres and a gassed power per unit volume of 0.5-10 kW. The defined correlation is ... 

Bioreactor scale-up remains a considerable challenge for producers of ERTs. One reason is that pivotal trials for LSDs are conducted on small numbers of patients, using quantities of enzyme that can be produced in smaller bioreactors, before the need to scale up quantities of product (and bioreactor size) to meet commercial demand. This is a conundrum that places considerable pressure on sponsors. The pitfall is that different scales may change the product profile and pharmacokinetics characteristics sufficiently that regulators require duplication of preclinical safety studies and even clinical trials after the pivotal phase 3 trials have been completed. This places an undue burden on sponsors seeking registration for these rare conditions. 

A general scheme for the development of an industrial process for alkaloid production is depicted in Pig. 1. On the basis of both fundamental research and feasibility studies the decision can be made whether an industrial production process is achievable. For the design of the process (production volume, process type, bioreactor size and type) detailed knowledge of both the kinetics of growth and product formation and physical properties (rheology, shear sensitivity) is essential. 

Optimum temperature is close to 20 °C and varies slightly with biocatalyst replacement policy. AC increases sharply over optimum temperature because of increasing cost of biocatalyst, since more frequent replacement is required as temperature increases. Below optimum temperature, AC increases smoothly due to an increase in energy consumption and in bioreactor size as temperature decreases below ambient. If no modulation effects are considered, AC is sensibly lower... 

Biological applications use smaller scale vessels because turbulence increases with scale, which leads to an increase in shear stresses as well. Airlift bioreactors have become popular in mammalian cell suspension apphcations for which shear stresses become important. Increasing the bioreactor size leads to an increase in the mechanical damage and lower cell densities (Martin and Vermette, 2005). 

The specific surface, a, is also relatively insensitive to the duid dynamics, especially in low viscosity broths. On the other hand, it is quite sensitive to the composition of the duid, especially to the presence of substances which inhibit coalescence. In the presence of coalescence inhibitors, the Sauter mean bubble size, is significantly smaller (24), and, especially in stirred bioreactors, bubbles very easily circulate with the broth. This leads to a large hold-up, ie, increased volume fraction of gas phase, 8. Sp, and a are all related... 

You have now demonstrated the capability of exponentially increasing microbial populations to rapidly produce protein. However, although such outputs of protein are possible in theory, they cannot be achieved in practice, since exponential growth cannot be maintained for such periods because bioreactors are limited in size. 

The submerged culture process continues to increase in terms of the percentage of dtric acid produced compared to that produced by the surface culture method. Tower bioreactors are preferred over stirred reactors because they cost less, there is less risk of contamination and they are less limited by size. 

A reasonable size of bioreactor, based on transport and handling considerations, is 200 m3, with a working volume of 150 m3. If file fermentation time is 48 hours and down time for reuse about 24 hours, then the total batch time is 72 hours. 

The mass transfer, KL-a for a continuous stirred tank bioreactor can be correlated by power input per unit volume, bubble size, which reflects the interfacial area and superficial gas velocity.3 6 The general form of the correlations for evaluating KL-a is defined as a polynomial equation given by (3.6.1). 

The power per unit volume is constant. From power consumptions in a bench-scale bioreactor, the necessary agitation rate is calculated for the scale up ratio, using Equation (13.2.1). The choice of criterion is dependent on what type of fermentation process has been studied. The following equation expresses relations for the impeller size and agitation rate in small and large bioreactors. 

The constant shear concept has been applied for bioreactor scale-up that utilises mycelia, where the fermentation process is shear sensitive and the broth is affected by shear rate of impeller tip velocity. For instance, in the production of novobicin, the yield of antibiotic production is dependent on impeller size and impeller tip velocity. 

Fig. 23. Average particle size dp after t = 120 h stirring for various impeller types and working conditions (left hand diagram data from [60]) and correlation with the maximum energy dissipation 8 (right hand diagram) stirred bioreactor with 4 baffles V = 6L D = 0.2m H/D = 0.96 zi=l...

For the same bioreactor mentioned above (operating at an impeller tip speed of 1.4 ms ), Dunlop et al. [57] predicted maximum Reynolds stresses in the impeller region of 32.4 Nm. However, if the energy is assumed to be uniformly dissipated throughout the vessel contents, then Eq. (5) will yield lower values. As calculated, the Reynolds stresses involve a length scale and the stress experienced by a particular entity will depend on its size. 

Gas-Liquid Mass Transfer. Gas-liquid mass transfer within the three-phase fluidized bed bioreactor is dependent on the interfacial area available for mass transfer, a the gas-liquid mass transfer coefficient, kx, and the driving force that results from the concentration difference between the bulk liquid and the bulk gas. The latter can be easily controlled by varying the inlet gas concentration. Because estimations of the interfacial area available for mass transfer depends on somewhat challenging measurements of bubble size and bubble size distribution, much of the research on increasing mass transfer rates has concentrated on increasing the overall mass transfer coefficient, kxa, though several studies look at the influence of various process conditions on the individual parameters. Typical values of kxa reported in the literature are listed in Table 19. 

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