An Economic Comparison of Three Cell Culture Techniques: Fed-Batch, Concentrated Fed-Batch, and Concentrated Perfusion
1. Janice A. C. Lim (Bioprocess Consultant) 1. John Bonham Carter (VP Sales)
2. Andrew Sinclair (Managing Director) 2. Jerry Shevitz (President)
Biopharm Services Ltd Refine Technology, LLC
英文职位名称Lancer Hou, East Street, 26 Chaplin Road, Unit 1107
Chesham HP5 1DG PO Box 691
United Kingdom Pine Brook, NJ 07058, USA
Contact: Andrew Sinclair Contact: John Bonham Carter
Tel: +44 (0) 1494 793 243 Tel: +1 793-993-3003
a.
participate
Photo 1: The operator is handling a ATF6 System in the plant.
Table 1: Guideline working volume sizes for each ATF system
lected nutrients during the growth culture cycle to improve productivity and growth. The culture is subquently harvested and the product recovered. Fed-batch culture has been an attractive choice for large-scale production due to its operational simplicity and familiarity as a carryover process from fermentation. However, fed-batch mode of operation typically also involves high start-up costs, resulting from the need for larger bioreactor plant capacity.
In perfusion culture, a continuous supply of fresh media is fed into the bioreactor while growth-inhibitory by-products are constantly removed. The increasing interest in the u of perfusion culture can be attributed to the higher product output from a reduced reactor size (hence, simplifying operation, cleaning and sterilization). The cell densities achieved in perfusion culture (30–100 x106 cells/mL) are typically higher than for fed-batch modes (5−25 x106 cells/mL) . The principal aspect of perfusion operation, which is different from fed-batch, is the added requirement of a cell retention device. Cell retention systems add a level of complexity to the process, requiring management, control and maintenance for successful operation. Perfusion bioreactors can suffer operational difficulties such as malfunction or failure of the cell paration device which can lead to shortening of the production run, leading further, to incread operating costs. This has previously limited their attractiveness.
In recent years, a new platform technology has been developed for biologicals production. The ATF™ System, introduced by Refine
Technology, LLC (Pine Brook, NJ). Ud in the alternating tangential flow mode, it is a low shear filtration system that inhibits filter-membrane fouling. This external cell-paration system is able to maintain
continuous culture for extended periods of time and offers the capability of rapid filter change without compromising the culture run.The ATF™ System allows incread volumetric productivity and reduced bioreactor size.吸血鬼日记第4季
Concentrated fed-batch and concentrated perfusion are two production techniques bad on the ATF™ System, which simultaneously nourishes the culture and concentrates the product within the bioreactor. The manufacturing methods permit great increas in cell and product concentrations as compared to fed-batch and perfusion. For example in the concentrated fed-batch production platform, one of Refine Technology’s pharmaceutical clients4 has reported a protein product titer of 17g/L with an unoptimized CHO cell process. Higher titers are expected as process
optimization continues. In the concentrated fed-batch operation, ultra high cell densities of (70–200)x
106 cells/mL have been achieved; similarly, extremely high cell densities in the region of (70–100)x106 cells/mL have been achieved in systems using the concentrated perfusion mode.
The system scales on a linear basis from 1L to greater than 1,000L and can be ud with traditional or disposable bioreactors and with all cell types including anchorage dependent lines. Table 1 indicates the working volume sizes for each ATF™ System in the scale-up process. The figures in the table are provided as guidelines. Actual capacity and vesl size depend upon process conditions.
This article compares the economic feasibility of a typical glycosylated protein manufactured using three production techniques – fed-batch (FB), concentrated fed-batch (CFB) and concentrated perfusion (CP). The Excel-bad process-cost modeling tool, BioSolve from BioPharm Services Ltd (Chesham, Buckinghamshire, UK), was ud for the economic asssment. The methodology, assumptions and key results of the cost model are described. The analysis will u the cost of goods (CoG) metric expresd on a per gram basis for comparability.
Evaluation Methodology
The commercial process-cost modeling software, BioSolve, is ud
to asss the process economics of the FB, CFB and CP production techniques. The CoG is a uful metric to quantify the operating costs of前途留学
a manufacturing option. BioSolve is an Excel-bad tool which determines the CoG by accounting for the indirect (fixed) overheads of the facility and the direct (variable) operating costs of the process. The fixed cost element consists of the capital charges, taxes, insurances while the direct costs include the consumables, materials, labour and waste management. Such an approach to estimate the CoG has been ud to evaluate the process economics of disposable membrane chromatography for u in biomanufacturing.
Figure 1 illustrates the basic BioSolve framework, which compris the ur interface, process definition, productivity and cost calculations, and model outputs. The interface allows a rapid asssment of the impact of pre-defined key input parameters (e.g. process scale, product titre, disposable options) on the model outputs. The core of the cost model consists of the process definition, which include the mass balances, equipment sizing, operating parameters and resource allocation. The calculated items include the plant productivity, labour requirements, consumable and materials usage, and equipment list. The costs of the components are maintained in a cost databa, which is built with data consisting of benchmarking information from veral biomanufactu
ring operations. The key outputs generated by the cost model are the facility throughput, bill of materials, CoG and capital investment. Ca Study Background
ncisThe model compares the manufacture of a glycosylated protein in three different production techniques – FB, CFB and CP. Figure 2 depicts the process quences for the production methods. The ed train for the FB process is longer as this option requires more steps to scale-up to the production cell culture bioreactor. The FB upstream process information was derived from commercially-relevant operations commonly practiced in the industry. The upstream information for the CFB and CP process was provided by Refine Technology, bad on data supplied by its clients. In the FB and CFB process, the removal of cells is achieved by centrifugation and depth filtration. Such recovery unit operations are not required in
the CP process as the cells are retained in the ATF™ System. The three process were assumed to have the same purification quence. This downstream quence was lected according to commercial relevance and sourced from details that are available in the public domain. Table 2
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summarizes the key operating parameters ud in the ba ca study.
Figure 1: Framework of BioSolve
Figure 2: Process quence (a) FB, (b) CFB and (c) CP for the production of a
glycosylated protein
For the FB process, all the ed steps included in the model are assumed to be batch fermentations. In the production cell culture step, two additional feeds are added to the bioreactor 3 days and 6 days after inoculation. The final growth pha takes about 13 days.
The CFB and CP ed quence is performed using a 200mL bag carried out in batch fermentation mode, followed by a 20L bioreactor also run in the CFB mode.
In the FB and CFB process, the recovery procedure is Centrifugation followed by Depth Filtration. A UF/DF step is ud to concentrate the
product stream prior to the Protein A unit operation. A Virus Inactivation step occurs. The product stream is subquently loaded onto an IEX Bind & Elute column and pasd through an IEX Flow Through column. A Viral Filtration step takes place prior to the UF/DF step. A final Sterile Filtration step is ud to purify the product. This purification quence is bad on typical unit operations for th
e production of a glycosylated protein, using typical solutions and timings. The overall yield for the downstream quence is 50% and 69% for the FB/CFB and CP process respectively.
a. Scenario 1 – Ba Ca木箱英语
Figure 3 capacity, which translates to lower initial investment costs. The CFB and CP technologies offer substantial capital charge reductions of 13% and 32% respectively relative to the FB process. The CoG per gram is inverly proportional to the annual product mass with the CP process producing the most product (265 kg/year) and FB producing the least amount of product (130 kg/year). This correlation between low CoG and high annual product mass is expected, as incread throughput of product will result in lower costs per gram produced.
In the CP process, the materials costs contribute a significant amount to the CoG. This can be attributed to the intensive usage of media over the extended perfusion culture cycle.
The model results indicate that the CP process (scenario 1c) is the most cost effective technology with the lowest overall CoG per gram of $87/g. The CFB process (scenario 1b) is slightly more expensive at $118/g. The FB process (scenario 1a) has the highest CoG ($158/g). The CFB and CP options provide cost savings of about 25% and 45% respectively when compared to the FB process.
Table 1: Guideline working volume sizes for each ATF system
(1) Cell concentration 50x106 cells/ml x 0.2L, diluted to 0.5x106 in 20L (2) Cell concentration 50x106 cells/ml x 20L, diluted to 1x106 in 1000L (3) vvd = volume of fresh medium/working volume of reactor/day.
(4)
Concentrated nutrients/factors are required during the run in the fed-batch process.
Figure 3: CoG results for the ba ca
新概念英语课件下载b. Scenario 2 – Changing the titre for the CFB process
The expression levels in the CFB process can vary depending on the type of cell line and duration of cell line development. In this ction, the effect of changing the product titre for the CFB process is examined while keeping the bioreactor size constant at 1,000L (Figure 4). In the worst-ca scenario, the titre is reduced from 17g/L (scenario 1b) to 10g/L (scenario 2a). In the best-ca scenario, the titre is incread to 25g/L (scenario 2b).
Reducing the titre, as expected, increas the CoG as the annual throughput is reduced. When the titre is reduced to 10g/L (scenario 2a), the CFB process appears less economical as the CoG ($184/
g) is higher
than the FB CoG at $158/g (scenario 1a). On the other hand, when the titre
is incread to 25g/L, the CoG reduces to $87/g. This improvement in the
expression level would make the CFB process much more attractive than the FB process with significant cost savings of 45%.
c. Scenario 3 – Changing the bioreactor size for the CP process因为单词
The bioreactor size ud for running perfusion culture mode can vary. This ction looks at changing the bioreactor size of the CP process while keeping the titre unchanged at 0.8g/L. Two other bioreactor sizes (500L and 750L) are considered in this analysis.
When the bioreactor volume is reduced to 500L (scenario 3a), the CP
process generate about the same annual output (132 kg) as the FB process (130 kg). The CoG of the CP process increas from the ba figure of $87/g (scenario 1c) to $133/g (scenario 3a). The CP process at 500L and 0.8g/L is still more economical than the FB process. It should be noted that t
he simplification of material and people flow provided by reduced bioreactor capacity in the CP facility cannot be calculated easily but may provide the key driver to such technology being increasingly ud in biomanufacturing.instruction是什么意思
Figure 4: CoG results for the FB process and CFB
process at 10g/L, 17g/L and 25g/L (1,000L bioreactor scale)
Figure 5: CoG results for the FB process and CP process using 500L, 750L and 1,000L (0.8g/L titre)
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