Increasing titers in mammalian cell culture has recently shifted the focus of process development efforts toward improving the economics of product recovery and purification processes. Some argue that a paradigm shift may be required to ensure timely and cost-effective delivery of future antibody candidates. This paper discusses the key process economic drivers and the impact of scale and titer on downstream processing (DSP) economic trade-offs. 

The success of therapeutic monoclonal antibodies (MAbs) in the treatment of indications such as cancer and autoimmune diseases has fuelled clinical trial activities; in recent years MAbs have become the fastest growing segment of the biopharmaceutical industry.

  Attractive returns coupled with the potential losses in revenue resulting from delays in product approval have made companies focus on speed-to-market rather than on improving process economics.

  However, MAbs are amongst the most expensive drugs—the annual cost per patient can reach $35,000 for antibodies that treat cancer conditions. With intensified competition on the way and increasing pressure from healthcare providers, the ultimate success of the next generation of MAb therapeutics will increasingly depend on economic factors.

  As a result, production costs and capacity use are becoming critical success factors for the industry.

The relative importance of key process parameters on the overall economic feasibility of a process varies with both the scale of operation and the titer assumed. As annual output and scale increase at a given titer or combined with titer increases, the relative importance of different cost categories are expected to change as follows (Figure 1): Overall cost of goods/gram (COG/g) decreases; direct/indirect costs: material costs rise considerably and dominate COG/g, labor and capital-dependent costs (overheads) represent a less significant proportion of COG/g; and upstream/downstream processing costs: downstream processing costs become a major component of COG/g.  

Figure 1.  Typical cost trends as scale increases for (a) total COG/g, (b) material, labor, and indirect (capital-related) costs, and (c) upstream and downstream processing operating costs. 

If the increase in annual output is also accompanied by titer increases, then the trends are expected to become even more pronounced. At small scales, fixed costs tend to dominate and any changes in raw materials will have minimal impact. However, as scale increases, the role of raw materials becomes more critical. It is therefore important to look at the distribution of the costs (upstream to downstream, and direct to indirect costs) and at the scale and titer of interest to prioritize optimization efforts. 

Efforts to lower COG/g must be targeted at decreasing the overall batch costs (e.g., reducing raw material costs) or increasing the overall productivity (e.g., increasing process yields). This section discusses specific downstream processing (DSP) drivers, which are of prime importance when handling multigram per liter titers.  

As titers increase to 10–15g/L, it is expected that this will have a profound effect on reducing COG/g, as long as the purification costs do not negate the cell culture gains. With increasing titers, the ratio of upstream to downstream costs shifts so that the downstream costs become more dominant. For example, Sommerfeld and Strube calculated that increasing the fermentation titer 10-fold from 0.1 to 1 g/L caused the ratio of upstream to downstream costs in their process to drop from 55:45 to 30:70. This shift reflects the fact that the upstream costs are inversely proportional to titer but the same is not true for the downstream processing costs.

  Increasing the titer to satisfy larger market demands increases the protein load on chromatography steps resulting in an increase in the number of cycles or additional investment in larger columns; this also produces larger volume loads on any subsequent filtration steps leading to longer filtration times or a need for larger areas.  

All these factors increase the downstream operating costs per batch. However, the overall COG/g can still fall if the increase in overall productivity outweighs the increase in downstream costs. Accordingly, as titers increase further, the downstream processing steps will become major contributors to the overall COG/g and offer greater potential for improvements and cost savings. Consequently, the downstream yield and material costs become significant cost drivers. DSP bottlenecks can lead to increased investment in larger equipment and longer batch durations. This results in increased running costs and decreased productivity and suboptimal COG/g values.  

The overall DSP yield is a function of the individual step yields and the number of downstream processing  steps, as has often been demonstrated using the plot in Figure 2. Improvements in step yields and the reductions in the number of steps have contributed to overall process yields increasing from 40 to 75% in recent years, with savings in cost of goods and investment, and allowing for higher facility throughputs.

Figure 2.  Overall yield as a function of individual step yields and the number of steps 

The impact of increasing step yields has been illustrated by Sommerfeld and Strube where increasing the average step yield in a seven-step downstream process from 85 to 95%, which increases the overall yield from ~30 to 70%, results in a 40% reduction in the downstream COG/g.

  Increasing step yields actually increases the equipment size or number of cycles required and, hence, the cost of the DSP steps, because each step needs to handle a larger load (chromatography) or volume (filtration). However, because more product is produced per batch, the COG/g typically falls with increasing yields. 

To maximize productivity and minimize investment and running costs, it is advisable to keep the number of downstream processing steps to a minimum.

  For antibody processes, this has encouraged the elimination of buffer exchange steps (diafiltration) that add little purification value by designing each chromatography step so that it can take the material eluted from the previous step where possible.

  Some companies have recently adopted processes that use only two chromatography steps while maintaining the desired purity levels.

  This requires an anion-exchange step to have additional selectivity to replace the intermediate purification and polishing steps. In cases where the contaminant profile, pH, and conductivity provide the opportunity to adopt this simpler process, time and cost savings can be achieved.  

In downstream processing, the distribution of raw material costs is highly dependent on whether or not resins and filters are reused. When treating these material components as disposables, resins and filters tend to dominate the material costs. Similar patterns can be seen in clinical manufacturing because these materials remain product specific. The impact of this operating strategy on the overall COG/g can depend on both the scale and the phase of development. The use of downstream consumables such as resins and membranes in a disposable fashion for a 200-L antibody process supplying early-phase clinical trials can provide both financial and operational savings.

  It has been reported that downstream processing using disposables can become a major disadvantage at the 10,000-L scale.

  This can be attributed to the fact that economies of scale result in a disproportionate effect on raw materials.

  Consequently, raw materials savings become more important for any process as the scale increases. 

The reuse of resins and filters involves a trade-off between reduced material costs and increased cleaning validation costs to determine the number of reuses with consistent performance. The higher the component cost or number of process steps, or the lower the validation costs, the greater the incentives to adopt filter or resin reuse.

  Chromatography resins, in particular Protein A, are often quoted as dominating purification raw material costs, owing to the high cost of the resin which is higher than ion-exchange resins. Large bioreactor scales of 10,000 L operating with a titer of 1 g/L, can result in Protein A resin costs of $4–5 million.

  Consequently, Protein A resins tend to be used in smaller quantities with multiple cycles, despite the complications of reuse validation.

  The reuse of Protein A resins can dramatically reduce their relative contribution to costs, making the costs associated with filters much more prevalent. In particular, virus filtration can represent a large contributor to purification material costs because of the costly membranes that are often used in a disposable fashion.

As the reuse of filters and resins increases, the cost of made-up buffers [chemical reagents and water for injection (WFI)] can account for a surprisingly large proportion of the costs, which can, in some cases, be greater than the cost of resins and filters. For example, at a fermentation scale of 20,000 L, approximately 140,000 L of buffer is required.

  The difficulty in estimating buffer costs partly reflects the large differences in the estimated costs of WFI, with values of approximately $0.20/L suggested for in-house generation and $3/L for vendor or contract manufacturing organization charges.

  Buffer costs have been quoted as varying between $2/L and $12/L.

  Efforts to reduce the volume and number of buffers required can lead to savings, because this naturally occurs when moving from a three-step chromatography platform to a two-step one. 

As increasing titers demand the use of larger downstream equipment or additional cycles, the requirement for buffers and WFI will also increase. In existing facilities, this can present retrofit challenges if demands exceed the capacity of the buffer preparation suites and WFI storage tanks or the rate of WFI generation. The use of buffer concentrates and in-line buffer mixing can help to reduce the tank size and floor space required for buffers.

The greater the mass load during chromatography, the greater the likelihood that the practical capacity of chromatography columns will be exceeded where the current limit of column diameters is 2 m.

  Under these circumstances, multiple cycles may be required and there is a risk that the downstream processing time will exceed the bioreactor time. This will reduce the potential throughput of the facility and impact on the COG/g. These large columns can also pose installation challenges in existing facilities if there is insufficient floor space and if they cannot fit through doors.

Process Economic Trade-offs for DSP Bottlenecks

Current efforts to avoid downstream processing becoming a bottleneck when handling larger masses include intensifying existing processes by enhancing capacity and speed. Such approaches can mitigate the need for extra investment in equipment and improve the process economics. However, with each of these approaches there are trade-offs and uncertainties that need to be evaluated to assess the impact on overall process economics.  

Chromatography Resin Dynamic Binding Capacity

Increasing resin binding capacity reduces column size requirements with concomitant drops in resin volumes required and buffer consumption [for equilibration, washing, elution, regeneration, and clean-in-place (CIP)] per batch. Novel resins have capacities that are more than twice those of the first generation resins.

  This can lower consumables costs which are a major component of COG/g at higher titers and demands,

  although the impact on the consumable costs will depend on the price differential between first and second generation resins. Increasing the binding capacity of affinity resins can have a greater influence on COG/g because affinity resins are more expensive than ion-exchange resins. With the cheaper ion-exchange resins, Sommerfeld and Strube highlight that a trade-off exists between the less pronounced drop in consumables cost and the increase in labor costs, which becomes more important at higher binding capacities, because of the longer processing times. However, given that most of the novel resins also allow higher flow rates, this may not be an issue. 

The first generation of cell-culture–based MAb processes used compressible chromatography resins. However, these impose severe limits on usable bed heights and flow rates when considering the expected increases in titer and scale.

  A move away from compressible resins and towards rigid resins that can handle higher flow rates reduces process cycle times and turnaround time and increases productivity.

  Flow rates with rigid resins can be three to five times faster than conventional compressible agarose resins.

  However, their use can also lead to increased buffer demands, higher pressure drops, packing complications, and shear stresses.

  New ion-exchange resins have recently been developed that can handle high flow rates (700 cm/h) and low back pressures (<3 bar). If the productivity increases outweigh the increased buffer demands, this will positively influence the COG/g.  

As mentioned earlier, resin and filter re-use can have a significant impact on the process economics at higher scales if the materials are expensive, despite the CIP and cleaning validation costs. Re-use also reduces the frequency of column packing which is time-consuming and costly when carried out on a large-scale.

  New resins are becoming available with increased stability when exposed to the harsh chemicals used for CIP; hence, they have longer cycle limits and can contribute to reducing the raw material costs.  

Platform processes provide a generic approach to antibody production that greatly reduces the development time while streamlining the regulatory aspects of processing. They represent a useful starting point for customization depending on the antibody being manufactured. Advances in resin properties have also allowed platform processes to emerge with two rather than three chromatography steps.

  This can help to alleviate DSP bottlenecks in existing facilities because a two-chromatography–step process occupies less floor space and consumes less buffer. Through such process intensification methods, Kelley predicted that a platform consisting of two chromatography steps with high capacity resins would be able to handle an annual output of 10 tons.

  Newer resins with the combined attributes of longer lifetimes, higher flow rates, and improved dynamic binding capacities will lead to improved platform processes for antibodies and contribute to significant reductions in downstream costs.

Research into alternatives for column chromatography focuses on methods that have the potential to effectively handle increased amounts of both the product and impurities (e.g., host cell proteins and antibody aggregates or isomers). Ideally, these alternatives should achieve a separation power equal to that of column chromatography while reducing the COG/g.

  When assessing the cost-effectiveness of these alternatives, it is important to consider not only the equipment sizes and resource consumption, but also the development and validation costs required.  

Membrane chromatography operating in flow-through mode is emerging as a popular alternative to anion-exchange chromatography steps in MAb purification,because of its rapid operation, ease of scale-up, and cost savings (Table 1).

  The dominant component in the distribution of raw material costs shifts from buffer costs in packed-bed chromatography to membrane costs; a membrane suitable for processing a batch of several thousand liters can cost several thousands of dollars and is disposable and not reusable. The key process economic trade-offs for anion-exchange applications therefore depends on whether the savings in buffer, labor, and overheads outweigh the high cost of the membranes. Critical variables that will affect the outcome of this cost comparison are the relative differences in the handling capacities assumed between anion-exchange membranes and resins, which dictate the sizes required, and the assumed WFI and buffer costs; higher values of these variables increase the economic attractiveness of membrane chromatography.

  The pace at which resin and membrane capacities improve will contribute to which operation secures its place in future platform processes. In cases where packed-bed and membrane chromatography offer similar COG/g, the real cost advantages may be in the development and validation costs that are significantly reduced with membrane chromatography because there is no column packing or cleaning validation.

Table 1.  Example of downstream process economic trade-offs 

As demand and titers continue to increase for MAbs, the DSP costs will become an increasingly dominant proportion of the COG/g with the DSP handling capacity representing a potential bottleneck that could reduce productivity. These factors have encouraged a shift in development efforts toward new DSP solutions that improve the process economics and alleviate bottlenecks. Consequently, the industry is taking advantage of improvements that affect the critical process economic drivers by looking to: improve the overall DSP yield and reduce the batch duration using platform processes based on two chromatography operations without intermediate buffer exchange steps; increase DSP capacity by taking advantage of improvements in chromatography resins that allow increased throughput over shorter times; and lower buffer demands and validation costs using new technologies such as membrane chromatography. 

These improvements will be important for facilities that already have large bioreactor capacities installed, and also for newer facilities that will probably be built with smaller capacities with flexibility in mind to allow rapid turnaround between campaigns for multiple products. Furthermore, if cheaper and faster expression technologies, such as glycoengineered 

, become more widespread, there will be an even greater spotlight on DSP costs. Although new DSP approaches may present complex and challenging problems to tackle, it is anticipated that this line of enquiry will dominate studies in the near future so that more cost-effective MAb platform technologies can evolve. However, with each of these new approaches there are trade-offs and potential risks that need to be evaluated to assess the impact on process economics. The capacity to cost such alternatives provides a common basis for such decision-making and will prove a vital tool for bioprocess designs in the future. Process economics can also be dramatically improved if the potency of MAbs is increased; recent efforts in this area are an encouraging sign for the future. 

This is an excerpt from the chapter entitled "Process Economic Drivers in Industrial Monoclonal Antibody Manufacture" in the 2009 John Wiley and Sons book Process Scale Purification of Antibodies edited by Uwe Gottschalk.

, is a senior lecturer (associate professor) at the Advanced Centre for Biochemical Engineering, Department of Biochemical Engineering, University College London, London, UK, +44 20 7679 4415, 

 1. Reichert JM, Rosensweig CJ, Faden LB, Dewitz MC. Monoclonal antibodies successes in the clinic. Nat Biotechnol. 2005;23:1073–1078. 

 2. Gottschalk U. The renaissance of protein purification. Supplement to BioPharm Int. 2006 June. 

Articles

Peer-Reviewed Research

The future of therapeutic MAbs lies in the development of economically feasible downstream processes.

Economic Drivers and Trade-Offs in Antibody Purification Processes

Most antibody manufacturers currently use a three column platform comprising Protein A affinity chromatographyfor product capture, followed by anion exchange (AEX) chromatography in flow-through mode to extract negatively charged contaminants, and then cation exchange (CEX) chromatography or hydrophobic interaction chromatography (HIC) in retention mode to remove positively charged contaminant species. This article presents a new process that uses membrane adsorbers for both flow-through and retention steps in antibody polishing. The process shows that a membrane adsorber in retention mode can be used efficiently in a commercial antibody manufacturing process.

Antibodies have been at the forefront of the biopharmaceuticals market for most of the last decade, and judging by the hundreds of antibody-based products at all stages of the clinical pipeline, these will continue to dominate the market for the foreseeable future.

  About half of all biopharmaceutical process trains have been designed for the production of monoclonal antibodies, and the industry has focused on a small number of platform technologies optimized to squeeze every last milligram of product from the feedstream.

  The vast majority of companies currently producing monoclonal antibodies (MAbs) use a three-column platform  approach comprising Protein A affinity chromatography for product capture, followed by anion exchange (AEX) chromatography in flow-through mode to extract negatively charged contaminants such as host cell protein (HCP), endotoxins, host DNA, and leached Protein A, and then cation exchange (CEX) chromatography or hydrophobic interaction chromatography (HIC) in retention mode to remove positively charged contaminant species including residual HCP and product aggregates. 

One disadvantage of this approach is that as cell culture titers increase, the downstream process becomes a bottleneck and costs increase in line with production scale, i.e., there is no economy of scale. The industry has been evaluating strategies to address such capacity and cost issues and has investigated different approaches, some at the low-technology end of the spectrum and others at the high-technology end. Among the latter, membrane adsorbers offer a flexible and cost-effective solution to some of the bottlenecks in antibody manufacture and are gaining acceptance as an alternative to traditional resin-based chromatography, particularly because they eliminate cleaning and validation costs, provide flexibility in production train design and scale-up, and can be used in viral clearance steps.

  Membrane adsorbers are thin, synthetic, microporous or macroporous membranes that are derivatized with functional groups akin to those on the equivalent resins. The membranes are stacked 10–15 layers deep in a comparatively small cartridge, generating a much smaller footprint than columns with a similar output. This reduces buffer consumption but increases the flow rate even though the bed height is much lower and there is a reduced pressure drop. Despite the increased flow rate, adsorption is efficient because the transport of solutes to their binding sites in a membrane adsorber occurs mainly by convection, while pore diffusion (the predominant mechanism in resins) is minimal. These benefits reduce process times to less than 10% of those associated with traditional stainless-steel columns.

  Another important advantage is the linear scale-up for parameters such as frontal surface area, bed volume, flow rate, and static binding capacity, while normalized dynamic binding capacity remains fairly constant at 10% or complete breakthrough.

  Scaling up process-scale operations therefore is a lot less troublesome than would be the case for standard resins, making the processing of up to 100-kg batches of antibodies a realistic proposal.

Currently, the benefits of membrane adsorbers are most apparent in flow-through applications because capacity constraints in retention mode, particularly at high loading rates, make resin chromatography a more attractive option for capture steps. Packed-bed chromatography certainly remains the preferred operation for capturing molecules of <200 kDa, especially when peak cutting and gradients are required for the separation of closely related species.

  However, for larger molecules (including most of the anticipated contaminants in antibody manufacture), membranes offer higher capacity and faster processing. For example, flow-through AEX for antibody polishing with a membrane adsorber can be conducted with a bed height of 4 mm at flow rates of more than 600 cm/h, providing a much higher frontal surface area to bed height ratio than is possible with columns. Even though the flow rate is much greater than would be possible with a column, there is sufficient retention time to reduce DNA, most HCP, and many viruses by up to four log reduction values, allowing membrane adsorbers to be used not only to separate the product from inert impurities but also as an integrated viral clearance step.

This article presents a new process developed at Philogen that uses membrane adsorbers for both flow-through and retention steps in antibody polishing. The set-up was used successfully with Teleukin, a new MAb fusion protein in Phase 1–2 clinical development at Philogen.

  After an initial capture step using a traditional Protein A column, the feed is loaded onto a Sartobind Q AEX membrane adsorber (Sartorius Stedim Biotech, Göttingen, Germany), which elutes directly into a Sartobind S CEX adsorber (Sartorius Stedim). The AEX adsorber is operated in flow-through mode to retain HCP, DNA, and leachate, while the CEX adsorber operates in retention mode, allowing the pure antibody to be separated from positively charged impurities. Because the eluate from the first adsorber is loaded directly and automatically onto the second, this can be regarded as a single, integrated polishing step that achieves up to 90% recovery and 99.9% purity as determined by size exclusion chromatography, cation exchange HPLC, and SDS–PAGE. 

The initial step in the purification train is a standard Protein A affinity chromatography separation carried out by loading the filtered culture medium from the bioreactor onto a column pre-equilibrated with phosphate-buffered saline (PBS). The medium is loaded overnight at room temperature under conditions that allow a contact time of approximately 30 s. The column is then washed through with an acetate saline buffer at pH 5.0, and the antibody is eluted with an acetate saline buffer at pH 3.0. The eluted product is adjusted to pH 6.0 and diluted to a final conductivity of 8.0 mSi/cm before loading onto the Q/S dual adsorber. As shown by the elution profile in Figure 1, it is possible to recover up to 90% of the product in this step at a >90% purity. 

Figure 1. Teleukin elution profile after the affinity chromatography step, based on absorbance at 280 nm 

The Polishing Step: A Q/S Dual Adsorber System

The polishing step is carried out using two Sartobind ion exchange adsorbers, a 7-mL Sartobind Q for the AEX step, and a 70-mL Sartobind S for the CEX step. These are arranged in series on an Amersham Akta Pilot (GE Healthcare, Uppsala, Sweden) as shown in Figure 2. However, for washing and elution, the two adsorbers can be operated independently, as shown in Figure 3. The two operational modes are selected by the operator using the Unicorn 4.12 management software (GE Healthcare). 

Figure 2. The Q/S dual adsorber system operating in series during the polishing step  

Before loading, the system is set to the "in series" mode and the adsorbers are washed through with 1 M NaCl and then sanitized in place with 1 M NaOH. After this treatment, the system is washed with 50 mM NaCl in pH 6.0 acetate/EDTA buffer (F2 buffer A). After washing and sanitizing, the eluate from the affinity purification can be loaded onto the Sartobind Q adsorber at a rate of up to 150 mL/min, a process that can take up to 40 min for a maximum 6 L of feed (or up to 500 mg of antibody), although greater amounts can be processed by splitting into multiple batches. In the "in series" mode, the flow-through containing the antibody elutes from the Sartobind Q step directly into the Sartobind S adsorber, wherein the target antibody binds to the sorbent. The system is then washed with 10 volumes of loading buffer before the operator switches it to the independent mode, allowing the Sartobind S adsorber to be washed with another 15 volumes of buffer at a flow rate of 350 mL/min before elution. 

Figure 3. The Q/S dual adsorption system operating in independent mode to facilitate elution from the Sartobind S adsorber. It is also possible to operate the Sartobind Q adsorber independently (not shown). 

Salt gradient elution is carried out in independent mode at 350 mL/min using a continuous gradient from F2 buffer A (50 mM NaCl, conductivity 8.0–8.5 mS/cm) to F2 buffer B (like buffer A but 1 mM NaCl, conductivity 86.0–87.0 mS/cm), which goes from 0% to 100% F2 buffer B in 20 volumes.  This step takes from 40 to 60 min and has a yield of up to 95%. A typical elution profile is shown in Figure 4.  

Figure 4.  Teleukin elution profile after the affinity chromatography step, based on absorbance at 280 nm 

The initial capture and purification step using Protein A affinity chromatography yields a product with >90% purity, but further polishing is required to reduce impurities and viruses to acceptable levels and conform with regulatory guidelines. The dual Sartobind Q/S disposable system tackles both positively and negatively charged impurities in a single step. The Q membrane is able to remove nearly 400 µg of DNA, 1.2 mg of Protein A, and nearly 30,000 EU of endotoxin (Sartobind Q5, surface area 5 cm

  equivalent to a 0.137 mL bed volume). Experiments to test the system for virus removal capacity are in progress. 

Table 1. Reproducibility and robustness of the dual Q/S polishing system, as determined by measuring yield and elution conductivity. Yield is expressed as [D]/[PD]%, where D is the amount of antibody dimer recovered after polishing and PD is the amount of antibody dimer submitted for processing. Elution conductivities (MC, mSi/cm) represent values between fraction T collection, and the value in parentheses represents conductivity at the edge of the chromatogram peak.

The reproducibility of these results has been tested by monitoring the yield (calculated as the amount of Teleukin obtained after polishing as a percentage of the input amount) and the elution conductivities of several runs. The results, as presented in Table 1, confirm the robustness of the process. The elution conductivities are highly reproducible both among lots and between runs, and the yield ranges from 82.2 to 99.9%, which shows that the process, when optimized, can recover nearly all of the input product. The Q/S dual adsorber system is therefore suitable as a manufacturing process for this product, and probably all MAbs. 

This final step in the process is a nanofiltration operation to mechanically remove endogenous or adventitious virus particles <20 nm in diameter. A 4-inch NFP Millipore cartridge filter (Millipore, Billerica, MA) is used for this step, and the efficiency of this device has been confirmed for a sample load up of up to 1.6 g of Teleukin in a volume of up to 6.5 L, loaded at a constant pressure (4 bar) using the Akta Pilot chromatographic system. The filtration flow rate is automatically adjusted by the Akta system to keep the pressure value constant. The flow rate usually ranges from 100 to 300 mL/min and sample recovery after filtration is usually very high (>90%).  

Several comparative evaluation studies have been carried out to determine the relative performance of membrane adsorbers and traditional packed-bed columns in flow-through mode, although direct comparisons can be difficult because different standards are used to measure process capacity and flow rate.

 When such issues are taken into account, the performance of Q adsorbers is encouraging, but until now there has been no similar evaluation of either Q or S adsorbers in retention mode. 

The process outlined above demonstrates that a membrane adsorber in retention mode can be used efficiently and robustly in a commercial antibody manufacturing process. The reproducibility analysis shows that the dual adsorber format operates successfully to remove both positively and negatively charged impurities from the feedstream, and because of the in-series linkage between the devices the dual system can be regarded as essentially a single, integrated polishing step. As membrane adsorbers become more widely accepted in antibody purification, it seems that not only flow-through but also retention chromatography steps in antibody polishing could be replaced by disposable membrane devices. Even so, the superior capacity of column chromatography for the capture of antibodies and other recombinant proteins means that resins are unlikely to be replaced for the foreseeable future in the primary capture step, which is the most significant bottleneck of all.

Leonardo Giovannoni, PhD, is the  head of CMC production and Marco Ventani is the head of  downstream processing, both at Philogen S.p.A., Sienna, Italy, +39 0577 588539, 

 Uwe Gottschalk, PhD, is vice president of purification technologies at Sartorius Stedim Biotech GmbH, Göttingen, Germany. Gottschalk is also a member of 

1.  Daemmrich A, Browden ME. A rising drug industry. Chem Eng News. 2005;Jun:28–42. 

2.  Gottschalk U. New and unknown challenges facing biomanufacturing. BioPharm Int. 2005; 18(3):24–28. 

3.  Thoemmes J, Kula MR. Membrane chromatography. An integrative concept in the downstream processing of proteins. Biotechnol Prog. 1995;11:357–367. 

4. Zhou J, Tressel T. Basic concepts in Q membrane chromatography for large-scale antibody production. Biotechnol Prog. 2006;22:341–349. 

A membrane adsorber in retention mode can be used efficiently in a commercial antibody manufacturing process.

Antibody Purification Using Membrane Adsorbers

Precipitation of process-derived impurities prior to capture chromatography in antibody purification offers a simple, robust, and economical method to efficiently remove Chinese hamster ovary (CHO) host cell proteins and DNA. By optimizing the major process parameters—pH, caprylic acid concentration, and mixing time—and understanding their interdependency, one can develop a cost-effective process step. When precipitation is applied directly in CHO cell culture, it combines the clarification and precipitation unit operations. The direct precipitation of contaminants results in seamless transition from upstream to downstream purification processes, particularly in high cell density and high titer cell culture. As a result, demand on the purification process is significantly lowered, and a simple two-step ion exchange process is sufficient to achieve therapeutic purity. 

Cation exchange (CEX) capture chromatography in non-affinity processes generally requires a feed conditioning step to lower the pH and conductivity in order to attain high binding capacity (~100 mg/mL). Typically, this is achieved by including a concentration and diafiltration step for primary recovery or, less preferably, through dilution and pH titration of the clarified cell culture bulk, resulting in large processing volumes.

Primary recovery TFF offers several advantages, including batch volume reduction, partial purification of process-derived DNA, and a cleaner feed stream for better column lifetime and performance.

  On the other hand, it is a lengthy unit operation, with high costs in terms of buffer consumption and TFF cassettes. Diafiltration also adds further constraints in the case of high titers by requiring even higher buffer and tank volumes at production scale.  

For those reasons, alternative technologies such as precipitation have been explored, particularly to achieve higher throughput for cell culture processes with high titers. Precipitation of the protein of interest has proven successful in food, blood, and enzyme manufacturing. For example, antibodies have been precipitated successfully at large scale by adding polymers of ethylene glycol (PEG)

  and research-scale precipitation of contaminants by charged polymers,

  cationic detergents, or short-chain fatty acids

  has shown promising application at the cell culture clarified bulk (CB) stage. 

In this study, we extend the application of precipitation that combines pH control and caprylic acid (a short-chain fatty acid) to remove process-derived impurities directly from CHO cell culture in the production of human monoclonal antibodies (HuMAbs). 

At low pH, the hydrophobicity of the octyl moiety of caprylic acid dominates and makes acidic proteins in the solution precipitate. Antibodies with basic pIs, however, have sufficient charge to counteract that hydrophobicity and remain in the supernatant. Thus, precipitation is carried out by first adjusting the pH to the appropriate level and then adding caprylic acid while mixing the contents. In this study, effective precipitation conditions were optimized primarily with clarified CHO cell culture supernatant and then extended directly to cell culture with limited development.  

To develop a robust and scalable contaminant precipitation step, three major parameters—pH, caprylic acid concentration, and mixing time—must be thoroughly studied. Caprylic acid concentration and pH are interdependent. Caprylic acid shows increasing efficiency in removing CHO host cell proteins (CHOP) and DNA as pH decreases. For example, cell culture harvest by precipitation at neutral pH does not remove sufficient levels of CHOP (Figure 1). However, controlling pH alone without adding caprylic acid does not remove significant amounts of CHOP; in the pH range of 4.0–7.0, CHOP levels were reduced by only ~20% in the absence of caprylic acid. The precipitation phenomenon exhibits two distinct phases of contaminant removal: first, a sharp (~2 log) decline of CHOP from pH 7.0 to pH 6.5, followed by steady 4-fold reduction to pH 4.0 (Figure 1).  

Figure 1. Optimizing pH for the precipitation of contaminants by caprylic acid in clarified bulk (caprylic acid 1%; mixing time 1 hour) 

The decision about whether to lower the pH further depends on the molecule's stability as well as the type of chromatography used in subsequent purification steps. In general, antibodies are less stable at lower pH. On the other hand, at higher pH, the binding capacity of the CEX column used in the following step will be significantly reduced. Therefore, conditions must be optimized to ensure the stability of the product and at the same time to maintain the high binding capacity of the resin. Caprylic acid concentration shows a threshold value for contaminant precipitation at any constant pH level (Figures 2a and 2b). Increasing the caprylic acid concentration from 0.1 to 0.5% at both pH 5.0 and 4.5 results in a ~2 log reduction of CHOP, followed by a negligible decline in CHOP precipitation. However, CHOP reduction was shown to be much more efficient at the lower pH level (4.5), even when the caprylic acid concentration was as low as 0.2% (Figure 2b). And at pH 4.5, product quality is not affected (see the stability discussion below). Therefore, in this case study, a caprylic acid concentration of 0.2 to 0.5% was considered effective.  

Figure 2. Optimizing caprylic acid concentration for the precipitation of contaminants in clarified bulk. 

Another process parameter that can influence precipitation efficiency is the length of time that the cell culture contents and the caprylic acid are mixed. Prolonging the mixing time from 30 to 120 min can significantly improve the removal of DNA and CHOP, with the longer mixing time needed particularly for the efficient removal of DNA (Figure 3).  

Figure 3.  Impact of mixing time on the efficiency of a precipitation step (pH 4.5, caprylic acid concentration 1%) 

Overall, these three parameters  (pH, caprylic acid concentration, and mixing time) must be tested interdependently. Other operating parameter ranges (e.g., temperature and mixing speed) also must be optimized for consistent performance of the precipitation step during scale-up. For the current studies, all the experiments at all scales were conducted at ambient temperature (20–25 °C). 

Direct Precipitation of Contaminants from Cell Culture: Two Case Studies

Caprylic acid also has been used successfully to remove host cell contaminants directly from cell culture. In two HuMAb case studies with CHO fed-batch cultures, caprylic acid precipitation resulted in significant reductions of CHOP and DNA contaminants. In HuMAb case study 1, CHO cell culture reached a peak cell density of 12 x 10

  cells/mL and a productivity of ~4 g/L. In HuMAb case study 2, cell densities were almost 10 fold (~120 x 10

  cells/mL) than in the HuMAb 1 study and the titer was also much higher: 14 g/L. 

The results of contaminant removal from cell culture in the first case study are shown in Figure 4. Precipitation with 1% caprylic acid with a mixing time of 1 h followed a pattern similar to that seen in clarified bulk precipitation (Figure 1). However, the only difference was a pH shift: cell culture required a pH level of 6.0 instead of pH 6.5 as in the case of clarified bulk to achieve the same level of CHOP reduction (Figures 1 and 4). This reduction also can be achieved simply by increasing the caprylic acid concentration, but  the cost of caprylic acid should be taken into consideration.  

Figure 4.  Optimization of pH for direct precipitation of contaminants from CHO cell culture (caprylic acid concentration 1%; mixing time 1 h) 

To precipitate CHOP from cell culture with a peak cell density of 120 x 10

  cells/mL, 1% caprylic acid at pH 4.5 with a mixing time of 2 h was tested. The CHOP level was significantly reduced, from 3.2 x 10

  to 12 ng/mg of antibody (Table 1). However, this experiment can be further optimized to define the most suitable pH level by keeping the caprylic acid level and mixing time constant. Based on the earlier results in cell culture, the right pH appears to be approximately pH 5.5.  

Table 1.  Scale-up of caprylic acid precipitation in clarified bulk and cell culture (pH 4.5 and 1% caprylic acid with a mixing time of 2 h) 

The target CHOP value to be achieved through precipitation also can be guided by purification process capability. Optimized CEX capture for case study 1 can lower the CHOP concentration from >1,000 ng/mg to <10 ng/mg. For example, if we aim to achieve a CHOP level of ~1,000 ng/mg, a pH of 5.5 may be sufficient for HuMAb 2, whereas for HuMAb 1, pH 6.0 will be optimum (Figure 4). If precipitation is performed at relatively low pH, it is important to monitor the stability of the filtered process intermediate after depth filtration (see the section on stability below). 

Caprylic acid precipitates form three layers after centrifugation. The top layer is a mixture of residual caprylic acid and the CHOP precipitate (Figure 5), and the bottom layer, the sediment, is enriched with precipitated DNA. Therefore, centrifugation may not be a practical option at large scale, in which case optimized depth filtration may be necessary.  

Figure 5.  Caprylic acid precipitation of contaminants  from clarified bulk  

Although depth filtration is commonly used to remove the biomass or contaminant precipitate, high cell density cell cultures pose several challenges for depth filtration. The significant costs involved in building filtration trains with multiple pre-filters makes a clarification step much more expensive than using a simple filtration assembly. By precipitating contaminants directly in cell culture, we can reduce costs by combining two processing steps for the recovery of product. In other words, the clarification of cells and removal of caprylic acid precipitate can be achieved by one depth filtration step. The removal of contaminant precipitates and scale-up are crucial for making caprylic acid precipitation a feasible procedure for large-scale manufacturing. In this study, caprylic acid precipitation has been successfully scaled up 2,000-fold in clarified bulk procedure and 200-fold in cell culture (Table 1).  

In a high density cell culture with continuous mixing mode with 1% caprylic acid, filtration capacity can reach ~80 L/m

. When the settlement of the precipitate is performed before filtration, capacity can be increased to as high as ~150 L/m

, which is almost as high as the filtration capacity of cell culture clarification. Another alternative is to add a filter aid to the mixture after precipitation is completed, followed by depth filtration. Filter aids such as Celpure (Advanced Minerals, Santa Barbra, CA) can help increase flux and filter capacity by separating the solids in the feed

  and increasing permeability. Using a filter aid before depth filtration for high solid contents, as in the case of caprylic acid precipitate, improves the quality of the feed stream as measured by turbidity units. The filtration capacity can exceed 1,000 L/m

  for bioreactor cell culture contents after precipitation with the addition of Celpure. Provided there is enough space to hold solids, filtration capacity is dependent on the filter pore size, filter aid amount, flow rate, and pressure.  

Integrating Precipitation into Non-Affinity Purification Schemes

Post caprylic acid precipitation, the filtered bulk is suitable to load onto either an affinity column or a non-affinity column. Here, we describe the integration of precipitation into a non-affinity purification scheme with CEX as the first chromatography step. 

The low pH condition after the caprylic acid precipitation procedure prepares the loading condition for high binding (~100 mg/mL) onto CEX resins. The levels of process-derived impurities have already been significantly reduced after precipitation under optimal conditions (Figure 6), leading to a much cleaner feed stream for the capture resin. As a result, the demand on the purification process is low enough that purification can be achieved by only two orthogonal separation steps. Typical step recoveries for precipitation are in the range of 80–99% for various HuMAb processes.  

Figure 6.  Product recovery, CHOP, and DNA levels following precipitation and CEX chromatography 

Impurities are at very low levels after CEX capture (<10 ng/mg HCP), making it possible to complete the purification process by adding just an anion exchange (AEX) membrane chromatography step (Figure 7a). The AEX chromatography in this case is mainly for adventitious virus removal;  a 3-log reduction of A-MuLV was still observed even at a conductivity level of 12 mS/cm (Figure 8), thereby accommodating the high conductivity CEX eluate with very low or no dilution.  

Figure 7.  Comparison of the tangential flow filtration (TFF)-based and precipitation-based process schemes 

With the most recent improvement in membrane chromatography, salt-tolerant membrane adsorbers have been developed to accommodate a wide range of salt concentrations in the load while still maintaining contaminant clearance and high flow rate characteristics. Using salt-tolerant interaction chromatography (STIC) disposable membrane modules (Sartorius Stedim Biotech, Goettingen, Germany), the process can be developed without any further dilution after CEX chromatography. STIC membrane chromatography was evaluated here for contaminant clearance to replace the AEX membrane step. The STIC membrane was able to significantly reduce CHOP from CEX eluates in a polishing step without any further dilutions (Figure 9). STIC can replace a Q membrane because the viral clearance capacity was reported to be 4–5 LRVs using the model bacteriophage virus ΦX174, especially at a high salt concentration where Q membranes have negligible clearance.

  It is clear that in a precipitation-based process where CHOP levels are extremely low after CEX, STIC may be able to process loads as high as those that Q membrane (20 g/mL) can when developed for viral removal only. Further studies are needed to confirm the scalability of STIC and its robustness to handle various buffer matrixes used in non-affinity processes.  

Figure 8.  Effect of salt concentration on viral clearance by  Sartobind Q at 20 g/mL load 

The residual caprylic acid concentration in the product was monitored for clearance during the purification process. In an experiment with 1% caprylic acid addition after depth filtration, only ~0.1% residual caprylic acid was present, which was further reduced to 0.003% after CEX chromatography alone. Several other fatty acids that are constituents of cell culture medium also were removed in a CEX flow-through fraction. CEX chromatography is a crucial purification step to remove cell cuture additives to significantly low levels in a two step purification scheme.  

Figure 9.  CHO protein (CHOP) removal profile for a salt-tolerant membrane used in flow-through mode (pH 7.0; conductivity 12 mS/cm). 

The precipitation step also significantly contributes to the viral removal strategy of non–Protein A purification schemes by adding ~4 LRV for A-MuLV.

  In our HuMAb case study 1, at pH 4.5 and with a mixing time of 2 h, 1% caprylic acid treatment provided a 4-log reduction in A-MuLV (Table 2). These studies will be extended for a range of model viruses in the future. The process capability of TFF-based and precipitation-based non-affinity purification processes to remove viruses is comparable when the CEX and Q membrane steps are performed under the same conditions for both process schemes (Table 2, Figures 7a and 7b).  

Table 2.  Viral clearance comparison for two non-affinity purification schemes 

The economics of two different HuMAb production processes, one using a TFF-based three-step non-affinity purification scheme and the other using a precipitation-based two-step process, were compared (Figures 7a and 7b) by modeling the processes with SuperPro Designer (Intelligen, Scotch Plains, NJ). The two models were compared for raw materials, consumables, and labor costs. After analyzing the processing time, overall process recovery, and cost (2 g/L, 5,000 L compared to 10 g/L, 2,000 L), it was obvious that the precipitation process scheme considerably lowers overall costs as a result of the efficiency of the precipitation step and because fewer processing steps are needed (Figure 10). As the product titer increases, these trends become more obvious, indicating that precipitation is favorable for high titer processes.  

Figure 10.  Economic analysis of precipitation-based non-affinity purification processes  

The quality of the product after caprylic acid precipitation and purification by CEX was compared with a reference standard that was purified by a TFF-based non-affinity process. The results of isoelectric focusing (IEF) gel analysis are shown in Figure 11. Furthermore, SDS–PAGE analysis, monomer purity assessment by SEC–HPLC, and activity by binding ELISA, are similar for the two products. The stability of caprylic acid treated filtrate at 2–8 °C was also monitored. The results of a case study showed that the filtrate can be stored at 2–8 °C for 34 d without any loss of activity (102% on day 50 of storage) and monomer purity. These stability studies will be continued on the final purified antibody to reveal the effect of caprylic acid precipitation process on long-term storage.  

Figure 11.  IEF analysis of stability samples following caprylic acid precipitation, stored at 2â8 Â°C 

The potential application of already successful and proven alternative technologies from other areas has been explored for antibody purification.

  High throughput process steps that are desirable for handling escalating titers will find their place in future large-scale manufacturing processes. Precipitation is one such alternative, and when applied to either the product or contaminants

  it significantly reduces demand on downstream steps. In addition, the precipitation of contaminants directly in the cell culture offers the further advantage of reducing unit operations at large-volume stages of the recovery processes. Because the precipitation step for cell culture can be performed directly in the bioreactor, using a disposable bioreactor is very convenient and lessens the burden of cleaning validation. Clearly, for a generic two-step non-affinity process, we have demonstrated a significant reduction in the cost of goods as well as improvements in productivity for a model HuMAb manufacturing process.

  However, economic comparisons need to be extended to various processes to account for different purification schemes and manufacturing scales used for each molecule. 

, is the assistant director of purification process development, 

 are scientists in purification process development, 

, is the assistant director of bioprocess development, and 

 is the senior director of process development,.all at Medarex, Inc., 908.479.2451, 

Precipitation prior to capture chromatography offers a simple, robust, and economical method to remove CHO host cell proteins and DNA.

Precipitation of Process-Derived Impurities in Non-Protein A Purification Schemes for Antibodies

Ion exchange–based membrane chromatography has already proven to be a powerful alternative to polishing columns in flow-through mode for contaminant removal. More recently, new membrane adsorbers have been designed for use in other applications. This article discusses a Sartobind phenyl membrane adsorber that has been developed for the manufacturing-scale production of biomolecules based on hydrophobic interaction chromatography (HIC) principles. The new adsorptive membrane combines the advantages of membrane chromatography with a high binding capacity for proteins comparable to that of conventional HIC resins. It represents a new tool for downstream processing, applicable for both bind-and-elute and flow-through operations.   

In the past decade, there has been a trend in the biopharmaceutical industry toward higher upstream productivity, i.e., higher cell culture titers, to meet the ever-growing clinical and commercial demand.

  This trend poses a challenge  for the industry—to develop and scale-up downstream purification processes capable of purifying multiple kilograms per batch. Although downstream processing faces significant bottlenecks because of heavy reliance on traditional bead-based chromatography.

  For example, although Protein A chromatography has undergone great optimization for throughput in the past five years, the expected rise in cell culture titers, combined with the need for larger masses of antibodies, indicates a need for further improvement of the existing Protein A resin to accommodate the demand for high-throughput production.

  Therefore, there has been significant interest in the development of new technologies that facilitate the processing of multikilogram batches at the lowest possible cost.

  Among the novel technologies developed, membrane chromatography has a high potential, particularly for process-scale monoclonal antibody (MAb) purification.  

Membranes  feature a more open structure than resins, so there is virtually no diffusion limitation, and as a result they offer many technical advantages that have been well described in the literature.

  Several applications have been discussed for MAb purification. Because of the generally high isoelectric points of MAbs, anion exchange (AEX) resins are typically used as a polishing step in a flow-through mode (i.e., the product does not bind while the trace impurities such as DNA, viruses, endotoxins, and host cell proteins are retained).

  Although the amount of impurities to bind is very low in such operations, conventional polishing scale-up leads to column oversizing because of the pressure and diffusion limitations associated with conventional bead-based chromatography. In contrast, a disposable membrane chromatography device has a convective mode of mass transport, which allows operation at significantly higher linear flow-rates (shorter residence times) than columns. As a result, disposable membrane chromatography device can have a much smaller volume than that of a conventional column when used in flow-through mode. This significantly reduces buffer consumption, processing time, and floor space requirements.

  Sartobind ion exchange capsules, for example, have been successfully implemented in various downstream processes for the removal of negatively charged contaminants.

  It is well documented in the literature that AEX membrane adsorbers are a powerful alternative to columns and can facilitate the development of new purification strategies for downstream processing. 

In addition to positively charged membranes, hydrophobic interaction membrane chromatography (HIC) has been described for the purification of a humanized MAb.

  Hydrophobic interaction is listed among the most commonly used chromatography methods for protein purification. For MAb purification, HIC is also often operated in flow-through mode.

  It is used to separate molecules based on their difference in hydrophobicity, and the technique has been presented as an efficient mode to remove dimers and high molecular weight (HMW) aggregates as a polishing step in a MAb purification process.

  In this article, we describe the use of an HIC membrane adsorber for the purification of a monoclonal IgG1 expressed using the PER.C6 human cell line. The hydrophobic membrane adsorber is based on hydrophilic regenerated stabilized cellulose with hydrophobic phenyl groups covalently attached to the base matrix. 

A Phenyl Membrane Adsorber for Large-Scale Purification

The Sartobind phenyl hydrophobic membrane adsorber was recently introduced to the market.

  It is assembled into 30-layer radial flow process capsules with an 8-mm bed height. The HIC membrane adsorber resulted from the combination of a newly designed macroporous membrane structure and covalently attached hydrophobic phenyl ligand. This membrane structure was designed for high flow rates and binding capacities. Furthermore, the hydrophilic stabilized regenerated cellulose membrane support possesses excellent mechanical and chemical robustness and demonstrates minimal unspecific interaction with proteins, even at high concentrations of a lyotropic salt.

  The functionalized membrane exhibits increasing binding capacity in increasing ammonium sulfate concentrations, as expected in HIC. Table 1 describes the binding capacity at 10% breakthrough for Sartobind phenyl challenged with three different model proteins.  

Table 1. Dynamic binding capacities (DBC) at 10% breakthrough for Sartobind phenyl membrane adsorber (3 mL membrane volume) challenged with three different model proteins. All proteins were prepared as 1 g/L solutions in 50 mM phosphate buffer at pH 7.0 with varying amounts of ammonium sulfate, and loading was done at 10 mL/min (18 s residence time). 

In a recent paper, the separation properties of the HIC membrane were evaluated and compared to those of a traditional HIC resin, phenyl Sepharose FF (low sub) from GE Healthcare (Chalfont St. Giles, UK).

  Using a protein mixture of cytochrome C, trypsinogen, and a polyclonal IgG, Sartobind phenyl demonstrated better resolution than the HIC resin under the experimental conditions used for the evaluation. In addition to the bind-and-elute operation, the authors demonstrated the application of Sartobind phenyl in flow-through mode for aggregate removal in a purification process for a recombinant protein. The protein was loaded with varying amounts of ammonium sulfate such that the aggregates bound while the monomer did not. After loading, the membrane was washed with deionized water to elute the bound aggregates. The aggregate removal increased with increased ammonium sulfate concentration, as shown in Figure 1. Protein yield was >95%, regardless of the salt conditions used in the loading buffer.  

Figure 1.  Flow-through mode operation with Sartobind phenyl membrane adsorber to remove high molecular weight (HMW) aggregates for the purification of an E. coli expressed protein. The figure shows the percentage of aggregates bound onto the membrane at increasing salt conditions. 

In this case study, a human MAb (IgG1, pI = 8.3, 150 kDa) was purified using the Sartobind phenyl membrane adsorber in flow-through mode. The protein of interest was produced at Percivia, LLC, using the PER.C6 cell line in a fed-batch process with a chemically defined growth medium. PER.C6 cells are E1-immortalized human embryonic retinal cells described in US patent 5,994,128.

  The crude media was clarified by centrifugation at 15,000

 followed by depth filtration and sterile filtration. The clarified media was partially purified by column chromatography. 

This material was loaded onto a 3 mL Sartobind phenyl nano membrane adsorber in 50 mM sodium phosphate buffer, pH 7.0, with ammonium sulfate concentrations ranging from 0.1 to 0.4 M. The product was collected in flow-through fractions and the membrane was then washed with equilibration buffer to recover the entire product. Samples of the flow-through fractions were assayed for aggregates by size exclusion chromatography.  

Figure 2.  Normalized aggregate levels of flow-through fractions collected while loading Sartobind phenyl membrane adsorber with partially purified antibody at various ammonium sulfate concentrations. (HMW: high molecular weight) 

Figure 2 shows that the binding of HMW aggregates improved with increased salt concentration in loading buffer. Although 100% breakthrough of aggregates was observed at 0.1 and 0.2 M (NH

, saturation was not reached at 0.3 or 0.4 M (NH

 when the membrane was loaded with up to 167 mg-MAb/mL-membrane. As shown in Figure 3, protein recovery was greater than 94%, even at the highest salt condition of 0.4 M (NH

, while aggregate reduction for the total flow-through pool was approximately 50%. In this study, a loading capacity of at least 167 mg-MAb/mL-membrane was achieved using the Sartobind phenyl membrane adsorber with yields of ≥94% and a final aggregate level <1% for a monoclonal IgG1.  

This alternative to column chromatography is suitable for flow-through as well as bind-and-elute purification operations.

BioPharm International-10-03-2009

Economic Drivers and Trade-Offs in Antibody Purification Processes

Antibody Purification Using Membrane Adsorbers

Precipitation of Process-Derived Impurities in Non-Protein A Purification Schemes for Antibodies

Hydrophobic Membrane Adsorbers for Large-Scale Downstream Processing

A Purification Platform for the Production of MAbs from Fermenters with Titers of 5 g/L and Beyond

Salt Tolerant Interaction Chromatography for Large-Scale Polishing with Convective Media

The Development and Application of a Monoclonal Antibody Purification Platform