Modern Biopharmaceuticals (Recent Success Stories) || Implementation of Advanced Technologies in Commercial Monoclonal Antibody Production

Part IVThe Rise of Monoclonal Antibodies – The Premium Class ofBiopharmaceuticals

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Modern Biopharmaceuticals: Recent Success Stories, First Edition. Edited by Jörg Knäblein.# 2013 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2013 by Wiley-VCH Verlag GmbH & Co. KGaA.

13Implementation of Advanced Technologies in CommercialMonoclonal Antibody ProductionJoe X. Zhou, Tim Tressel, Xiaoming Yang, and Thomas Seewoester

13.1Part I: Commercial Antibody Process Development

13.1.1Introduction

Antibodies have become the most successful molecule modality in the developmentof modern biological medicine in the past decade. Monoclonal antibodies in factdominate current biotechnology drug pipelines [1–3].The commercialization of these products (Table 13.1) has led to significant

advancement in the knowledge of antibody process and production in recent years.The economic commercial manufacturing of a therapeutic antibody requires a highyielding, robust, and operationally cost-effective process that delivers a product withhigh purity [4]. In this chapter, we discuss current unit operations of a commercialantibody manufacturing process with specific considerations to critical parameters.Approaches to improve production cost of goods (COG), including reduction ofbuffer/water consumption, decrease in manufacturing time, and enhancement ofviral clearance are described.

13.1.1.1 Essential Considerations for a Commercial Process DevelopmentThe most important target of commercial process development and optimization isto minimize production COG by increasing product throughput while maintainingproduct quality. To do so, unit operations must be optimized toward maximumproductivity. More effective technologies or techniques that were not available atinitial development should be considered and applied during the development of acommercial process. A comprehensive characterization of operational parametersand critical raw materials must be conducted to support a robust commercialprocess. In cases of upstream and downstream process changes, a combination ofsingle-variable studies and design of experiment (DOE)-based studies are consid-ered essential to determine acceptable ranges for critical operational parameters.

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Modern Biopharmaceuticals: Recent Success Stories, First Edition. Edited by Jörg Knäblein.# 2013 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2013 by Wiley-VCH Verlag GmbH & Co. KGaA.

For commercial process development, product comparability is the ultimate test.The product purity and impurity profiles, biological activity, and product stability areexpected to be comparable to the product made in the early phase process.

13.1.1.2 Major Challenges for Upstream and Downstream ProcessesThe major challenges for upstream and downstream processes can be classified intwo levels: overall challenge for the process and specific challenges for each unitoperation.Process robustness is one of the largest overall challenges for any process and

special attention is needed during the upstream and downstream process optimi-zation. To provide flexibility during the commercial manufacturing, the processneeds to be robust across multiple production sites and variations in key processparameters. Thus, it is important that the ranges determined during the subsequentprocess characterization or bench scale validation need to be broader than theacceptable operation ranges (Figure 13.1).The cell culture process for antibodies contributes significantly to the initial

productivity and often to key product quality attributes. The variability in controlparameters such as temperature, pH, pCO2, and dissolved oxygen (DO) levels

Table 13.1 Antibody therapeutics in drug market.

Antibodytherapeutics

Companies Launchtime

OKT-3 Ortho/J&J 1986ReoPro Centoco/J&J/Eli Lilly 1994Rituxan Idec Pharm, Genentech, and Roche Ltd 1997Zenapax Roche Inc. and PDL 1997Herceptin Genentech and Roche, Ltd 1998Remicade Centoco/J&J 1998Simulect Norvatis 1998Synagis MedImmune and Abbott Laboratories Ltd 1998Mylotarg Celltech and Wyeth 2000Campath ILEX Pharm., Millennium Pharm. Inc., and Berlex

Laboratories, Ltd2001

Zevalin Idec Pharm. and Schering AG 2002HUMIRA/TRUDEXA

CATand Abbott Laboratories, Ltd 2002

Bexxar Corixa Corp. and GSK 2003Raptiva Genentech 2003Xolair Novartis Pharma AG, Genentech, and Tanox, Inc. 2003Erbitux(cetuximab)

ImClone, Pfizer, and Onyx 2004

Avastin Genentech 2004Tysabri Biogen Idec Inc. 2004Vectibix Amgen Inc. 2006Soliris Alexion 2007

252j 13 Implementation of Advanced Technologies in Commercial Monoclonal Antibody Production

should be minimized to obtain consistent performance parameters such as peaktotal cell density, viable cell density, titer, and harvest yield.The downstream process should have a broad operation range so that it can accept

certain variability from the upstream process without affecting product quality andyield. Operational parameters, such as pH, conductivity, and wash and elutionvolumes, should be tested outside buffer specifications (optimization or acceptableranges in Figure 13.1) to guarantee process robustness. With a robust process,consistent yield and comparable product attributes are achievable across manufac-turing sites.Unique challenges in upstream process development start with the comparison of

themaster and the working cell bank (WCB) to ensure a consistent startingmaterial,followed bymedium and process optimization to increase yield. The scalability froma bench scale bioreactor to a commercial scale bioreactor of 10 000-l ormore needs tobe fully understood. Consistent and controlled performance across all scales iscritical. All raw materials should be assessed toward variability and potential impacton product quality attributes.A larger challenge and often less explored field is the cell separation step, often

performed through continuous centrifugation and depth filtration. The scalabilityand performance of the centrifuge are very difficult to address and predict asdifferences are seen across centrifuge venders, models, and scales. Tangential flowfiltration (TFF) technology is currently not considered the first choice for monoclo-nal antibody harvest because the subsequent step, protein A, is a very robust unitoperation for a direct capture. This affinity chromatography-based capture/purifi-cation unit operation can remove most impurities including proteases and achieveover 98% purity. Selection and optimization of the capture column is critical in theharvest step, because this step may determine the risk level of process and productimpurities. The risk level in a non-affinity-based unit operation can be significantlyhigher for a purification process.

Buffer SpecRange

Center Point

Alert Limit

Action LimitEdge of Failure

Range Identification of Operation Parameters

Opt/AcceptableRange

CharacterizationRange

FMEA(Worst Scenario Testing)

Range

Figure 13.1 Process parameter setup in different ranges during optimization andcharacterization.

13.1 Part I: Commercial Antibody Process Development j253

The purification process follows with polishing and viral filtration steps. Theirfunction is to remove aggregates (AGG), host cell proteins (HCP), DNA, and leachedprotein A, and to also ensure a sufficient viral clearance power. The challenges hereare how to reduce manufacturing time and cost, while meeting process and productquality criteria.

13.1.1.3 Dosage and Bulk Product PurityAt present, most monoclonal antibodies on the market are IgG1 and IgG2, withapparent molecular weight of 150 kDa. Different monoclonal antibodies may havedifferent solubility profiles, but in general a solubility level of >100mg/ml can beexpected [5].Current literature indicates that impurity levels in antibody therapeuticsmay be as

low as<5 ppm for HCP,<10 ng/dose for DNA,<0.5% for dimer or aggregates, and<5 ppm for leached protein A in the final bulk drug substance (BDS) [6]. However,not all monoclonal antibody therapeutics on the market may reach such low levels ofimpurities.

13.1.1.4 High-Titer Cell Culture Processes and its Impact to Downstream ProcessesTo meet the market demands, companies generally strive toward a high-titer cellculture process [7]. But a high-titer cell culture process in turn may also create arelatively higher level of product-related impurities including AGG and dimericforms. In addition, monoclonal antibodies are complicated glycoproteins subject toglycosylation heterogeneity and other modifications such as N-terminal pyrogluta-mate and C-terminal lysine variants, methionine oxidation, asparagine deamidation,disulfide bond scrambling, and fragmentation. Most monoclonal antibody down-stream process platforms in industry using protein A affinity as a major capture stepfollowed by polishing steps can not remove these different variants to acceptablelevels. This challenge has to be overcome through a smart cell clone selection. Thus,clone selection is extremely critical and sufficient time should be taken to select thebest clone. This strategy often prevents the costly situation of having to change to amore suitable clone during late-stage process development.

13.1.1.5 Viral Clearance StrategyMammalian cell culture systems are susceptible to potential contamination byadventitious viruses that may be introduced through raw materials or failures inprocess controls [8]. In addition, noninfectious retrovirus-like particles are producedby Chinese hamster ovary (CHO) cells, which are consistently monitored andquantitated by electron microscopy [9–11]. A major functional expectation for alldownstream polishing steps is to remove viruses in addition to separation of traceamount of process and product-related impurities. A common optimization strategyfor the downstream process is to maximize product purity from the first capture/purification unit, while building adequate viral clearance capability in the polishingsteps. This also can be challenging because commonly used chromatography stepshave a low selectivity to capture different viral particles with variable genotypes, pIs,and sizes [12].


13.1.2Upstream Process

Recombinant mammalian cells are the expression system of choice to producemonoclonal antibodies. CHO andmurinemyeloma (NS0) cell lines are the two linesusedmost widely in the industry [8]. A typical cell culture process starts with thawinga frozen vial of a working cell bank (WCB), and the cell population is expandedthrough a series of seed trains in different culture vessels. The culture is thentransferred to a production bioreactor where the cells continue to grow and secreteproduct into the culture broth. To achieve high product titer, high cell mass and cellviability need to bemaintained in the production bioreactor. Two of themost popularprocess modes employed in industry are the fed-batch process (feeding concen-trated nutrient solutions to a batch production bioreactor) and the perfusion process(maintaining and recycling the cells to a bioreactor while replacing old media withcontinuous freshmedium supply). The fed-batch process is often used because of itsscalability, ease of operation, reduced contamination risk, and high volumetricproductivity. An illustration of the unit operations for a typical upstream process isshown in Figure 13.2.Although stainless steel bioreactors are still the common choice for large-scale

production, disposable bioreactor systems have become available. Disposable bagsystems from various manufacturers are now commonly used during seed cultureexpansion [13]. Compared to traditional bioreactors, disposable reactors have theadvantage of not needing any costly cleaning or a validated clean-in-place (CIP) orsteam-in-place (SIP) procedure. Overall, disposable reactors can often be more costeffective than stainless steel bioreactors [14].Cell culture process optimization is an integrated activity involving cell line

selection, medium development, and bioreactor condition optimization. High titers

Figure 13.2 A typical cell culture process composes vial thaw, seed expansion, and productionstages. Source: Reprinted with permission from Ref. [4].


of 5–10 g/l and cell densities of >20 million cells/ml have been recently reported infed-batch cultures [15,16]. A cell-specific productivity of>20 pg/cell per day can nowbe routinely achieved for production cell lines [17] and sometimes even up to 100 pg/cell per day have been reported. Enhancement of specific productivity per cell isaccomplished not only by the selection of highly productive cell lines, but also byoptimization of medium compositions and bioreactor operation conditions.

13.1.2.1 Cell Line DevelopmentCell line development starts with transfection of the gene of interest into the hostcell, subsequent gene amplification, and final cell clone selection. Amaster cell bank(MCB) is generated to establish a cell seed stock for a reproducible manufacturingprocess. All these steps are conducted during early phase development. Theprocedures from clone selection to MCB are described in our previous publication[4]. At the beginning of the commercial process development phase, a working cellbank (WCB) is created from the MCB. In case, the early clone is inadequate to meetcommercial productivity or product quality targets, a new cell line might bedeveloped. Recent advances in high-throughput clone screening and selectiontechnologies allow to effectively identify highly productive clones right from thebeginning. Screening large numbers of different cell clones with high expressionand secretion levels can be achieved using a fluorescent-tagged antibody against theproduct expressed on the cell surface or secreted in a microgel bead, followed by afluorescence-activated cell sorting [18,19]. Another approach involves miniaturizedbioreactors or shake flasks that can simulate standard production bioreactorconditions, including nutrient feeding [20,21]. All these approaches usually employsome level of automation technology to streamline the liquid handling and analyticalcapability.

13.1.2.2 Media and Feeding Strategy DevelopmentRecent significant antibody titer increase in cell culture can be attributed greatly toimproved cell culture media. In general, media development for a fed-batch processinvolves batch medium and feed concentrate as well as feeding strategy optimiza-tion. Optimization of cell culture processes is often regarded as cell line dependentand must be based on the metabolism and nutrient consumption by a specific cellline. Several approaches can be used systematically for medium optimization, suchas single-component titration, spent medium analysis, and medium blending [22].Because of regulatory concern about serum and other animal-derived components,the addition of plant-based or other nonanimal derived hydrolysates to a chemicallydefined base medium is a common approach to increase cell density, cultureviability, and productivity. Hydrolysates are protein digests composed of aminoacids, small peptides, carbohydrates, vitamins, and minerals. In addition, hydroly-sate peptides can also act as growth factors and production stimulators. Nonanimal-derived hydrolysates from soy, wheat, and yeast are commonly used in cell culturemedia and feeds. However, because of its composition complexity and sometimesseen lot-to-lot variability, hydrolysate can be a significant source of medium andprocess performance variability.


Alternatively, to hydrolysates additions, a more concentrated chemically definedbase or feedmediummay be used. For such an approach, it is critical to remove saltsas much as possible in order to avoid high osmolality, which will negatively impactthe cells. Optimization of any feeding strategy needs to consider nutrient consump-tion, by-product accumulation, osmolality, and the balance between cell growth andproduct secretion [23].

13.1.2.3 Bioreactor Process and ControlThe most common commercial scale for bioreactors is between 10 000 and 20 000 l[24,25]. Advancements that enabled very large-scale bioreactor processes are mainlydue to (i) the understanding of homogeneous, low-shear mixing and (ii) theimproved controllability of gas transfers, in particular, the impact of pCO2 levelson cells.In large-scale bioreactor cell culture, dissolved oxygen (DO) is maintained at a

moderate level through direct sparging of pure oxygen gas. Recent understandingalso points to the critical role of CO2 concentration in scale-up [25]. It was reportedthat a low DO level resulted in a decreased glycosylation of antibody N-glycan chains[26]. Therefore, appropriate gassing and mixing strategy is considered critical for asuccessful scale-up.pH is considered as another key parameter. Higher pH (�7.0) may promote initial

cell growth, whereas slightly lower pH may facilitate antibody production [27].However, high pH condition could increase anaerobic metabolism, with moreconversion of glucose to lactate. High lactate accumulation in turn may causeexcessive base addition and result in high osmolality. The combined condition mayreduce cell growth or accelerate cell death.Maintaining the optimal temperature is another critical process parameter to

consider. Sometimes changes in temperature can induce desired effects as it canaffect the cell cycle in CHO cell cultures. A temperature shift from 37 �C to 30 �Cafter 48 h postinoculation can retain cells in the G1 phase longer and therefore delaythe onset of apoptosis [28].With regulatory agencies promoting process analytical technology (PAT) and

quality by design (QbD) approaches to ease the regulatory review process, it isexpected that future major advancements in bioreactor process controls will includeeffective on-line monitoring instrumentations for the measurement of cell density,CO2, osmolality, metabolites, and protein concentrations. This will enable a moredirect and comprehensive process and product quality control in real time.

13.1.2.4 Impact of Cell Culture Process on Product Quality AttributesAs mentioned previously, monoclonal antibodies are complicated glycoproteinssubject to glycosylation heterogeneity and other modifications such as N-terminalpyroglutamate and C-terminal lysine variants, methionine oxidation, asparaginedeamidation, disulfide bond scrambling, aggregation, and fragmentation. Cloneselection and cell culture conditions have the potential to change these productquality attributes and may impact product efficacy. During any process optimizationit is critical to monitor product quality changes at every development stage.


Glycosylation variation, which may impact IgG in vivo functions and productstability [29], is one of the most sensitive quality-related attributes. Most mouse-derived cell lines are known to add Gal-a1, 3-Gal onto the antibody heavy chain N-glycan. A high-throughput screening approach can be used to select low Gal-a1,3-Gal clones in the early process development stage [30].The effects of cell culturemedia and conditions on antibody glycosylation have been

extensively studied [31,32]. Glucose, ammonia, dissolved oxygen, dissolved CO2, andculture osmolality have been reported to cause glycosylation changes in different celllines. These factors affect the activity of monosaccharide transferases and/or sugartransport to Golgi, which is the major glycosylation site in mammalian cells.Although the integrity of antibody backbone is usually unchanged in different cell

lines and culture conditions, some modifications can happen during cell cultureprocesses. The IgG heavy chain C-terminal is conserved as lysine, which is normallytruncated through posttranscriptional modification. Presence or absence C-terminallysine can result in product charge variants. Carboxypeptidase B specifically cleavesC-terminal lysine residues. Medium trace element concentrations can affect theactivity of these metal enzymes.

13.1.3Downstream Process

Antibody downstream process has advanced significantly and continues to improveat a rapid pace. Specific technology for each unit operation is discussed in thefollowing sections.

13.1.3.1 Harvest and Capture Process

13.1.3.1.1 Harvest and Recovery The harvest process separates the antibodyproduct in the culture broth (conditioned medium) from the cells. A typical harvestdiagram is illustrated in Figure 13.3 [33,34]. Continuous centrifugation is commonly

Figure 13.3 A typical harvest diagram for antibody process. Source: Reprinted with permissionfrom Ref. [4].


used to harvest large-scale cell culture broth ranging from 500 to 15 000 l to generatecell-free supernatant. To generate a supernatant with low cell turbidity, choices of fullor partial shot, total number of shots, and buffer volume per shot during thecontinuous centrifugation process need to be optimized.

13.1.3.1.2 Supernatant Clarification The supernatant is then further clarifiedthrough a depth filtration followed by a 0.2-mm absolute filtration [35]. Althougha sterilizing grade filter is used to prevent potential contamination, sterility is notclaimed for this step. For depth filtration, selection of filter type, flux, and totalmembrane surface area needs to be evaluated for an optimal performance. A well-selected filter can also remove up to 50% of DNA and 15% of host cell protein at theneutral pH. The filtrate is then loaded on a protein A column [34].

13.1.3.1.3 Protein A Affinity Chromatography Protein A affinity process is currentlythe predominant capture column choice in industry because of its high flow rate (upto 500 cm/h), excellent binding capacity (25–50 g/l resin under the loading condi-tions at a pH 6–8 and conductivity 10–>20mS/cm), superior large column packingprofile, and extensive removal of process impurities such as HCP, DNA, mostpigments, and media components. Protein A resin also has the ability to removesome critical impurities such as proteases; therefore, it minimizes the potential riskof clipping by proteases.Use of protein A resin is also considered a virus removal step (Table 13.2) [36].

Additionally, incubation at a low pH of 3.4–3.8 at room temperature for 45–90min,which follows a low pH buffer elution from protein A resin, has been demonstratedto efficiently inactivate enveloped viruses such as X-murine leukemia virus (X-MuLV, Table 13.3).

Table 13.2 Viral removal by protein A resins.

Resins MuLV load MuLV LRV MMV load MMV LRV

Agarose-based 7.34 2.77 8.47 2.39Glass-based 7.07 3.31 8.16 2.62

Source: Reprinted with permission from Ref. [36].

Table 13.3 Viral inactivation by a low pH incubation post protein A column elution.

MuLV viral inactivation MMV viral inactivation

pH 3.7 3.7Incubation time (min) 60 60Viral load (LRV) 6.75 7.5Log reduction value �3.75 �1.5



There are two major protein A resins available: Agarose-based by GE Healthcare(Piscataway, NJ) and glass-based by Millipore (Billerica, MA). Both resins are robustenough to handle a high flow rate with acceptable chemical resistance, includingresistance to high-concentration urea, guanidine-HCl (GuHCl), and reducingagents. Dynamic binding capacity ranges from 25 to 50 g of antibody per liter ofresin. Both resins can handle up to 200 cycles [34].Because of different physical properties, each resin has its advantages and

disadvantages. Agarose-based resins have low nonspecific binding; however, theyhave a disadvantage of high operational backpressure for a large packed column (IDof 45 cm and above) with a bed height>20 cm. Tomaintain the operational pressurebelow 30 psi, a flow rate below 400 cm/h is used. Alternatively, the bed height can bereduced while maintaining the same flow rate to achieve a desired operationalbackpressure [34].Glass-based resin can handle higher flow rates up to 700 cm/h with fewer

backpressure issues. However, the resin is sensitive to caustic solution and has ahigher nonspecific binding property. Solvent-containing buffer wash can be used toreduce HCP to a sufficiently low level. Several organic compounds have been testedand developed as efficient wash reagents, including tetramethylammonium chloride(TMAC) and tetraethylammonium chloride (TEAC); however, these reagents aretoxic and costly [37]. A few less toxic, more efficient, and economically acceptablereagents recently have been developed as wash reagents [37].The mechanism for reducing nonspecific binding for glass-bead resin is not well

understood. One theory is that the dual action of solvent/ionic effects disrupts theassociation between HCP and silica [34]. Another theory is that the nonspecificbinding sites are partially coated by these reagents inhibiting HCP nonspecificbinding; however, no published data are available. Table 13.4 lists HCP levels fromglass-based resins with or without wash compared with the level from Agarose-based resin [34].Normally, multiple cycles on protein A column is routinely used to process each

cell culture lot. The resin is stripped between cycles and the column is regeneratedonce the last cycle is completed. When culture titer is low (50–200mg/l), particularlyfor an initial phase study, it makes sense to use a short residence time with a highflow rate to shorten loading time. As the antibody titers increase, a long residencetime is preferred to achieve a high binding capacity. Thus, the number of cycles can

Table 13.4 A comparison of HCP and DNA removal on protein A resins.

Protein A load Glass-based resin Agarose-based resin

No wash Wash No wash

Titer (mg/ml) 1.15 9.1 8.45 6.69HCP (ppm) 410 943 8367 1275 2496DNA (ppm) 13 054 830 4396 3077 6876



be reduced [34,38], and high product throughput and significant buffer volumesavings can be achieved.A loading temperature in the range of 15 �C� 3 �C is recommended not only for

the protection of product but also of protein A resin itself. Process contaminates,such as proteases, will digest protein A ligand from the resin quickly at roomtemperature, leading to a high level of leachable protein A in the pool [39].In summary, the following critical process parameters should be optimized:

residence time for optimized binding capacity; wash solution composition, pH,and wash volumes to ensure low levels of HCP and DNA in the pool; and elutionbuffer composition and pH to reduce pool turbidity, dimer, aggregate levels, poolsize, and conductivity for the next step [34,40]. Optimization should always bematched with large-scale capacity in mind.

13.1.3.1.4 Low pH Viral Inactivation This step follows protein A affinity purifica-tion. A low pH elution is needed to remove and collect purified monoclonalantibodies from the protein A affinity resin. The pH of the elution buffer solutionis typically 3.0–3.4, and the pH of the protein A elution pool is 3.6–4.2, depending onthe ionic strength of the elution buffer. Because protein A pool can be easily titrateddown to a pH of 3.4–3.8 depending on the stability of monoclonal antibodymolecules, a low pH viral inactivation method is commonly implemented immedi-ately following protein A chromatography. The monoclonal antibody pool is incu-bated at the low pH for 45–90min to inactivate potential contaminating viruses. Thismethod is efficient at 20 �C and above for inactivating enveloped viruses like X-MuLV. A less-efficient inactivation is expected at temperatures of 15 �C or below. Incontrast, low pH is not effective in inactivating nonenveloped DNA viruses such asMMV at any temperature (Table 13.3) [36].The kinetics of inactivation of MMV and X-MuLV at pH 3.7 and ambient

temperature are illustrated in our previous publication [12]. In these studies,maximum virus inactivation for both viruses was reached after 20min of incuba-tion. As expected, an insignificant 1- to 1.25-log reduction in virus load titer wasobtained for MMV. In contrast, a >3.5 log reduction in virus load titer wasachieved for X-MuLV particles. The efficiency of X-MuLV inactivation duringthe low-pH step can be significantly reduced if the pool solution contains aminoacids such as glycine [36].In general, low pH (3–3.5) is required for monoclonal antibody elution in protein

A chromatography. Consequently, this harsh conditionmight alter biological activityof the monoclonal antibody or cause protein aggregation for low pH-sensitivemonoclonal antibodies. As a result, mild elution conditions at pH values higherthan 4.5 have been studied. For example, several pseudo-affinity resins have beendeveloped that use protein elution at a high pH. Once the new elution methods areadopted for monoclonal antibody production, the low pH viral inactivation unitoperation may no longer be as favorable for viral inactivation.Innovative technology for viral clearance is needed. Among these methods in

development, ultraviolet irradiation technology (UVC) is perhaps the applicationof choice [41–44]. Although this technology has demonstrated that the high log


reductions for several viruses are achievable (Table 13.5), some limitations, such asdenaturation of monoclonal antibodies and suitability for large-scale application,need to be fully investigated.

13.1.3.1.5 Protein A Pool Filtration After low-pH viral inactivation, the pH of thepool is titrated up to a range of pH 5–8 for the next step. The pool may become turbidwhen the pool pH is raised above 3.7. The degree of turbidity varies frommolecule tomolecule. Therefore, filtration using either a nominal depth filter or a 0.2-mmabsolute filter is required. A well-selected depth filter can not only remove turbiditybut also reduce HCP and DNA [34]. Although a viral clearance function of depthfiltration has been demonstrated, qualification of the depth filter for processconsistency remains a challenge [45,46].In summary, optimization parameters that need to be considered in protein Apool

filtration are the depth filter selection, surface area required, and operation flux rate.In addition, a successful process output cannot be achieved if the filter isundersized.

13.1.3.2 Polishing ChromatographyTwo or three additional chromatographic steps, called polishing unit operations, arerequired to achieve desirable product purity. These steps may function for (i) high-molecular weight aggregate reduction, (ii) trace host cell protein (HCP) clearance,(iii) clip/isomer isolation, (iv) DNA reduction, (v) leached protein A clearance, and(vi) viral clearance. The challenges can vary significantly frommolecule to molecule.Several chromatography unit operations are available to address these challenges [5].

13.1.3.2.1 Cation Exchange Chromatography Cation exchange (CEX) chromatog-raphy in bind-and-elution manner has proven to be an extremely powerful step toremove product-related impurities that protein A affinity cannot. CEX resins arescreened based on HCP removal, high product dynamic binding capacity at arelatively high conductivity, and separation power to remove target protein variants.For an ideal resin, about 70–80% of HCP and most of DNA/RNA and endotoxin arein the flow-through (FT) fraction. Deamidated or acidic species are usually separated

Table 13.5 Viral inactivation power with UVC device.

Viruses X-MuLV PRV Reo-3 BVDV HAV Monoclonal antibodyrecovery (%)

Enveloped Yes No No Yes No —

Genome type RNA DNA RNA RNA RNA —

LRV at 50 J/m2 4.3 6.5 6.8 5.7 4.7 97LRV at 38 J/m2 3.3 3.2 2.6 3.8 2.9 98LRV at 30 J/m2 3.3 3.4 2.0 3.4 3.0 98

Source: Data from Gottschalk U, Reif O-W, Tarrach K, Mora J, Kaiser K. An orthogonal concept forreliable virus clearance in biomanufacturing. Presented at: Downstreamprocessing Technology Forum;November 2, 2004; London, UK.


in the front of the peak and amidated or basic species and dimers/aggregates can beisolated in the post peak [35].Elution can be accomplished in either stepwise or linear gradient manner. Linear

gradient elution can provide better purity, process control, process monitoring,reproducibility, and PAT conformance [38,47]. In contrast, stepwise elution hasproven to be mechanically simpler with higher product pool concentration [47].Recently, a linear pH-conductivity hybrid gradient elutionmodewithin a narrowpH

range has been developed [48,49]. Compared with salt gradient elution, pH gradientelution provides equivalent product purity with a higher yield, a smaller pool volume(up to 50% of volume by salt gradient), and low pool conductivity. Zhou et al. [49]demonstrated that pH gradient CEX chromatography can be successfully scaled upwith a recovery of>95%. Aweak cation resin is ideal for pH gradient elution owing toits narrow titration curve. Several pairs of buffers can be potentially used for variousnarrow pH gradients for purification of antibodies with different pI values.

13.1.3.2.2 Hydrophobic Interaction Chromatography Hydrophobic interactionchromatography (HIC) is a powerful tool to remove dimers and aggregates inbind-and-elution manner. However, it uses a high concentration of salts, producesrelatively low yield, and generates a low separation power for other product-relatedisomers. Therefore, HIC in bind-and-elution manner is growing less popular inantibody production. Instead, flow-through (FT)-HIC has been developed to removeaggregates with a relatively high yield. However, a slow flow rate is needed, becauseaggregate binding on HIC resins depends on residence time.

13.1.3.2.3 Ceramic Hydroxyapatite Chromatography Ceramic hydroxyapatite(CHT/HA) chromatography with a sodium phosphate gradient has also beenused as a robust step to remove dimers, aggregates, and leached protein A[50,51]. Gagnon et al. [51] recently demonstrated the powerful application of asodium chloride linear gradient using human antibody IgG4 in CHT chromatogra-phy. Compared with the partial aggregate separation with 30 column volumes (CV)of linear gradient from 5 to 300mM sodium phosphate at pH 6.5, the monomerpeak was completely separated from the aggregate peak with a 30 CV linear gradientfrom 0 to 1MNaCl in 5mM sodium phosphate at pH 6.5. The successful separationleft the majority of DNA, endotoxin, and leached protein A in the 0.5M phosphatestrip pool. The purity attributes for the pools by the two different elution solutionsare summarized in Table 13.6 [51]. Indeed, CHT column operation in the bind/

Table 13.6 Purification summary of IgG4 using NaCl versus phosphate in CHT chromatography.

Endotoxin (EU/ml) DNA (ng/ml) Protein A (ng/ml)

Starting material >500 1194 19.96NaCl pool <0.05 0.67 0Phosphate pool 16.6 3.8 1.27

Source: Adapted with permission from Ref. [51].


elution mode can be used as a powerful polishing step for large-scale monoclonalantibody production; however, several issues, such as resin lot-to-lot viability, resinlife, and viral clearance capacity, still remain.

13.1.3.2.4 Anion Exchange Chromatography Anion exchange (AEX) is perhaps themost powerful tool to remove a variety of viruses, DNA, and endotoxins. In mostcases, AEX chromatography is carried out using flow through (FT) mode, with theimpurities binding and the product of interest flowing through. To prevent a processbottleneck, the use of conventional packed-bed chromatography with FT-AEXrequires columns of very large diameter to permit high volumetric flow rates[35]. As proper flow distribution in production columns requires minimum beddepths, bed volume becomes significantly large [52]. As a result, such packedcolumns are dramatically under-used because they are packed for speed and notoptimized for binding capacity. The limitations of AEX columns have led to thedevelopment and use of membrane chromatography (MC) or membrane absorbers.Current membrane chromatography products offer a convenient alternative to resinchromatography in the purification of monoclonal antibodies [53,54]. In addition toQ membrane development, HIC and S membranes might also have potentialapplication as disposable modes and are discussed in Part II of this chapter.

13.1.3.3 Viral FiltrationAt present, size-based nanometer filtration technology is perhaps the most robustviral removal unit operation. Nanometer filters can be divided into two classes: 50-and 20-nm pore sizes. Large-pore filters are efficient in retaining large-particleviruses like X-MuLV and pseudorabies virus (PRV). On the other hand, filters with asmall pore size (20 nm) remove both large viruses and small virus particles such asMMV and Reo-3. Additional information on filter performance and issues is foundin our previous publication [45].For optimization, the following parameters need to be tested:

1) Feed pH and protein concentration. Usually, feed pH is more important thanprotein concentration. When the pH for viral filtration is close to protein pI, theprotein solution possesses a high viscosity that will generate a high operationalpressure, leading to filter clogging [55].The relationship between viscosity and protein concentration is not linear. A

protein solution of 10 g/l has a higher viscosity than a solution concentration at5 g/l; however, the viscosity values are not significantly different. When theprotein solution has a concentration >10 g/l, the viscosity difference is consid-erable. For example, monoclonal antibody solutions at 8.5, 15, and 20 g/l haveviscosities of 1.00, 1.02, and 1.05 cP, respectively, determined at 20 �C.

2) Viral filters from different vendors. Different vendors provide different filters. Thus,Vmax and flux studies are important to select the best filter for the process.Usually, the Vmax value using a scale-down model determines the filter capacityand flux determines the process time. Through these studies, a cost-effectivefiltration unit operation can be generated.


3) Protein solubility in different buffers. Monoclonal antibodies show differentsolubility profiles in different buffer systems. Therefore, a solubility studymust be done before viral filtration studies.

13.1.3.4 Aseptic FiltrationFor more than 30 years, aseptic filtration with 0.2-mm membranes has remainedvirtually unchanged because of the quality, predictability, and flexibility of thetechnology.Producing monoclonal antibodies at high concentrations (up to 150 g/l) showed

that aseptic filtration can be challenging since a limited volume of antibodytherapeutics at such concentrations will readily clog most 0.2-mm membranes. Astep recovery of only 80–85% is normally obtained with the regular 0.2-mmmembranes. This leads to a significant loss of recovery for the entire process. Toovercome the challenges, filter vendors newly designed several membrane configu-rations (Table 13.7) [34].Our recent studies were performed using a purified antibody solution at 150 g/l to

evaluate new sterilizing grade filters with two different constant operating pres-sures: 5 and 15 psi. Figure 13.4 presents the Vmax values for the sterile filter testedusing the vendor scale-down models. Among the filters tested, the dual-layer filterswere the best performers (Figure 13.4). Based on the Vmax values obtained from thescale-down experiments, dual layer filters used in our aseptic filtration produceda consistent recovery >95%. In one of our recent runs, only 1% was lost when a4800-ml bulk was processed using a 500 dual layer autoclaved sterile filter. In contrast,about 15–20% was lost by clogging and filter exchange during the process when thesingle-layer filter was used.Newly designed aseptic filters with a high process capacity are not without

problems. Currently, the major dilemma with the aseptic membrane is nonspecificbinding of Polysorbate-20 [56] and the levels of extractables. Based on our recentdiscovery, all the filters tested remove a certain amount of Polysorbate-20 duringfiltration. Polysorbate-20 is currently widely used for prevention of aggregateformation. Thus, solutions to overcome these issues are critical [56].

Table 13.7 Sterile filter membrane comparison.

SuporEKV

SuporLife SartoporeII

ExpressSHF

ExpressSHC

Durapore

Membranematerials

PESa) PES PES PES PES PVDFb)

Layers 2 2 2 1 2 1Pore size (mm) 0.5/0.2 0.8/0.2 0.45/0.2 0.2 0.5/0.2 0.2

a) PES: polyethersulfone.b) PVDF: polyvinylidene fluoride.Source: Data from Zhou J, Tressel T, Hong T, Solamo F, Dermawan S. Utilize Current SeparationTechnology to Achieve a Cost-Effective Robust Process for Recombinant Monoclonal AntibodyProduction. Presented at: Antibody World Summit; July 24 to 27, 2005, 2005; Jersey City, NJ.


For measuring extractables, the sterile filter membrane industry uses similarmethods as for water for injection (WFI) [57]. The newly developed total organiccarbon (TOC) assay can detect organic impurities at the ppb level. In addition, aconductivity assay detects both the inorganic and organic impurities in the bulkedWFI. Both methods are more accurate for quantitative and reliable measurementsthan the methods used previously.The extractables in general are generated during heat-based autoclave or

c-irradiation. Based on the certificate of analysis provided from the vendors, after1- to 5-l water flush for each 1000 device, the level of extractables is low enough to be ofno safety concern. Therefore, for a large volume of purified bulk monoclonalantibodies, aseptic membranes are not problematic.

13.2Part II: Implementation of Membrane Technology in Antibody Large-ScalePurification

13.2.1Introduction

Several chromatography modes have proven very useful in the removal of traceimpurities and virus [75–77]. Among these modes, flow-through anion exchange(FT-AEX) is perhaps the most powerful tool to remove a variety of viruses [58], DNA[59,60], and endotoxin [61]. Near neutral pH at low conductivity (3–7 mS/cm), manyviruses, DNA, endotoxin, and a large percentage of host cell proteins are negativelycharged and will bind to the AEX resin, whereas the typically basic (i.e., positively

0

300

500

600

Vmax

(l/m

2 )

100

400

200

5-psig15-psig

0

300

500

600

100

400

200

5-psig15-psig

Figure 13.4 Process capacity (Vmax) evaluationof the aseptic filters available commercially at amonoclonal antibody concentration of 150 g/l,pH 5.0. Source: Data from Zhou J, Tressel T,Hong T, Solamo F, Dermawan S. Utilize Current

Separation Technology to Achieve a Cost-Effective Robust Process for RecombinantMonoclonal Antibody Production. Presented at:Antibody World Summit; July 24 to 27, 2005,2005; Jersey City, NJ.


charged at neutral pH) antibody species will not. FT-AEX is therefore commonlyused as a polishing step in Mab production [59,62].However, the use of conventional packed bed chromatography with FT-AEX

requires columns of very large diameter to permit high volumetric flow rates inorder to prevent a process bottleneck [6,62]. Proper flow distribution in productioncolumns requires a minimum bed height to prevent nonuniform packing andinadequate header design [63]. The required bed height leads to a significantly largebed volume. Consequently, such packed columns are dramatically oversized becausethey are designed for speed and not for binding capacity. Unfortunately, this designfor speed greatly increases the cost and complexity of the AEX unit operation. Theselimitations of AEX columns have led to the development and use of membranechromatography (MC) or membrane adsorber (MA).

13.2.1.1 Pros and Cons of Using Q Membrane Chromatography as a Purification UnitQ and other charged MA devices have been in development for chromatographypurposes for more than 15 years [63–65]. Some limitations for these devices used forproduction scale are (i) distorted or poor inlet flow distribution [66–68]; (ii)nonidentical membrane pore size distribution [69,70]; (iii) uneven membranethickness [70]; and (iv) lower binding capacity [66,70]. The first three weaknessescan be improved to some degree when multiple-layer configurations are used in thecurrent MA device [70]. This configuration for the Q membrane is used in viralvaccine production [71] and DNApurification for gene therapeutic agent production[72–76]. Q anion exchange adsorber devices have also been used for endotoxinremoval at process scale [77]. Low binding capacity is still a major disadvantage in abind/elute protein purification mode [70]. The low binding capacity is attributed tolower surface-to-bed volume ratio as well as flow distribution problems, and this isextremely difficult to overcome [66,70,78].In a flow-through mode, however, the limitations and weaknesses are no longer

considered as major issues, particularly, when FT-MA is used as the polishing stepfor antibody purification. FT-MA chromatography is useful to remove impuritiesbelow 1% concentration, including viruses [34,62,79]. The demand for a flow-through high-throughput polishing step in large scale provides a great opportunityfor the application of Qmembrane chromatography. In this concept, the current MAoffers many advantages over packed bed resin chromatography [80–82]. It has goodviral clearance with fast convective flow. The maximum linear flow rate reported forQ-SepharoseTM Fast Flow is about 200 cm/h [83]; while linear flow rates of 600 cm/hare achievable, based on our data [84]. The large pore size with theMAprovides highbinding capacity for large biomolecules such as viruses and DNA at a higher flowrate compared with columns. Antibody recovery in the flow-through mode iscomparable to columns, normally about 98–100%. Buffer usage for MA can bereduced to only 5% of that for a conventional packed bed chromatography.

13.2.1.2 Historical Studies of Q Membrane Chromatography in Antibody ProductionA thorough evaluation of a 10-layer Sartobind1Qmembrane scale-downmodel wascarried out by Knudsen et al. [62] from Genentech, Inc in 2001. The study reported

13.2 Part II: Implementation of Membrane Technology in Antibody Large-Scale Purification j267

that the Sartobind1Qprovides a 2.3-log reduction ofmurine leukemia virus (MuLV)at a flow rate of 620 cm/h and at a capacity of 2000 g/l sorbent or about 550 g/m2. Theviral log reduction decreased with increased antibody load. For example, 1.0 log,0.5 log, and 0.3 log of viral reduction (log reduction value, LRV) were obtained atcapacities of 1100, 1650, and 2755 g/m2, respectively. In contrast, Q-SepharoseTM

Fast Flow provided a viral clearance power of >5.1 LRV at a capacity of 50 g/l. Thepoor viral removal data with the Sartobind Q membrane was likely due to its 10-layered format instead of the 15-layer device used in our recent studies.Gallher and Fowler [82] from Millennium Pharmaceuticals, Inc. used the Sarto-

bind1 Q capsule module (Q-10 inch in 15 layer format) in large-scale antibodyproduction for viral removal. The capacity was estimated as 450 g/m2 with�4.0 LRVfor respiratory enteric orphan III (Reo-3), MuLV, and pseudorabies virus (PRV). In2004, Pieracci et al. [85] from Biogen Idec, Inc. presented an antibody process withMustangTMQat 4mS/cm,pH8.0, andfluid velocity of 40MV/minas the last polishingstep.The load capacitywas estimated tobe587g/m2or 2.7 g/ml ofmembrane volume.An antibody recovery of about 100% was reported for three lots of 6.7, 5.6, and 3.2 kg,respectively. The monitored outputs were antibody recovery (90–99%), membranepressuredrop (5–9psig), andhost cell protein removal (12–23 ppm in theflow throughpool). Viral clearance of 5.4–6.5 logs for MuLV was reported.At BioProduction 2004 conference, Zhang et al. [86] from Abgenix, reported a study

on a Sartobind1 Q used in an antibody pilot plant run to produce material for atoxicology study. The scale-down model with Sartobind1 Q75 at neutral pH andconductivity below 3mS/cm demonstrated excellent viral removal: 5.57 LRV forMuLV, 7.28 LRV for Reo-3, 6.77 LRV for MVM, and 5.67 LRV for PRV [87]. However,by estimation; the process capacity was limited to be 480 g/m2 or 1750 g/l mem-brane volume.At the SCI Membrane Chromatography Conference in 2004, Jerold Martin [60]

presented data generated with the Pall MustangTM Q, confirming many of theadvantages of membrane chromatography. Process capacity information was notreported, so the economics for the process could not be determined. TheMustangTM

Q membrane provided >8.1 logs of DNA removal and good viral clearance withMuLVand PRV. Poor results were obtained for REO andMMVunder the conditionstried.

13.2.1.3 Operation Units for Membrane ChromatographyLinear velocity (cm/h) is recommended as a common flow rate unit in membranechromatography [34,62,79]. Other units such as membrane volume/min or devicevolume/min have been used as flow rate units in the membrane chromatographyliteratures. However, it is difficult to use membrane volume/min or device volume/min to compare the real flux for the membranes of different thickness. For example,Sartobind1 Q is 3–5mm in pore size, 0.0275 cm in thickness with a 15-layer formatdevice, and MustangTM Q is 0.8 mm in pore size, 0.01375 cm in thickness with a 16-layer format device.Historically, a membrane volume was used to estimate the membrane’s capacity

for impurities removal, similar to that used for columns. Now we believe that it is


more accurate to compare capacities using total surface area in a flow-throughmodebecause membrane devices are made of different materials with different layers,pore sizes, and thicknesses. In addition, total surface area is easily used to comparethe performance of membranes from vendors. Furthermore, the membrane pricingfor different module sizes is determined based on a total surface area used for aprocess. Either grams of antibody processed per liter of membrane or grams ofantibody processed per meter-square of membrane may be used as membraneprocess capacity units. Grams antibody processed per meter-square of membrane(g/m2) is our preferred capacity unit to compare Q membrane efficiency.

13.2.2Analysis of Q Membrane Scale-Down Models and Cost Factors in Large-ScaleProduction

Cost saving has been the major argument between membrane and resin vendors;therefore, it has to be analyzed accurately. A general cost estimate was made basedon a 15 000 bioreactor with a 1 g/l titer, and a process yield of 90% (leads to a Qload of 13 500 g product). This model case is extrapolated from data obtained fromseveral 2000 l-scale runs. Table 13.8 provides a comparison of buffer consumptionand manufacturing time by Q column and Q membrane. As shown in Table 13.8,FT-MA can save up to 95% of buffer and 66% of process time. The costcomparison is shown in Table 13.9 with a WFI cost of $3.00 per liter. The processusing membrane is about $54 268 at a loading capacity of 500 g/m2. The processcapacity of Q SepharoseTM Fast Flow commonly used for antibody purification isabout 100 g/l. Thus, at $800/l, the resin initial cost is about $138 000 for 172.5 l ofresin, which would be packed into a 100-cm column to a bed height of 20 cm. TheQ resin can be reused for at least 200 cycles in our calculation, although, ingeneral, the resin is discarded after 100 cycles due to the cost of other items suchas column cleaning-in-place and storage, column packing/qualification, andproduct changeover. Thus, the resin cost per cycle is only $690. To process13 500 g of antibody, the cost of the process using QSFF resin is about $18 067under these conditions. Thus, it appears that the Q column unit operation is morecost effective than the Q membrane at a process capacity of 500 g/m2.A process capacity of 2000 g/m2 or 2 kg/m2 would be required to make economicsense with Q-MA in a large-scale process (Table 13.9). At this capacity, the total costfor one cycle, including the Qmembrane ($12 420), equilibration, and wash buffers,is about $15 560. Thismakes theMA very competitive. The cost of any unit operationis based upon the plant running the process.

13.2.2.1 Q Membrane Process CapacityThe main purpose of the AEX step is efficient viral reduction. The typical feedmaterial has total impurities, such as DNA andHCP at concentration of<1%. Thus,data from the viral clearance scale-down model should be used to determine themembrane’s maximum capacity. Scale-down studies with defined operating param-eters (flow rate, buffers, feed stock, and operational temperature) are used to


determine this capacity. The use of at least two viruses is suggested for membranecapacity determination. Ideally, one virus would be a small size, such as the nakedDNA virus minute virus of mice (MVM) with a 20-nm particle size and anotherwould be larger, such as the enveloped virus MuLV with a 100-nm particle size[53,54,79].

13.2.2.2 Issues of Current Scale-Down ModelsCurrently, most scale-down models including Sartobind1 Q75 are not suitable in aviral clearance study due to process capacity limitations generated by operationalbackpressure. Scale-down models do not have a liquid flow path similar to thatfound in the large-scale capsule and thus tend to generate extremely high opera-tional backpressure when a high flow rate is applied (�450 cm/h). The scale-downmodel, such as Sartobind1 Q75, is made with stacks of several flat sheets. In the

Table 13.8 Example of buffer use and processing time for Q-packed bed chromatography and Qmembrane chromatography.a)

Items Q membraneb)

(DV �5.5 l)Time(h)

150 l columnc) Time(h)

Setup — 1 — 1WFI flush — — — 0.5Prime buffer lines/installfilters and instruments

— 0.5 — 0.5

Prime skid and filters — 0.5 — 0.5Preequilibration — 0 3 CV¼ 450 l 0.47Equilibration 10 DV¼ 55 l 0.04 5 CV¼ 900 l 1Loading 2700 l 1.95 2700 l 2.87Wash 10 DV¼ 55 l 0.04 3 CV¼ 450 l 0.47Regeneration NA 0 3 CV¼ 450 l 8

Slow flow rate and longcontact time for thoroughcolumn cleaning

Storage NA 0 3 CV¼ 450 l 0.47Setup to clean skid — 0.5 — 0.5Clean skid — 0.5 — 0.5Storage-skid — 0.5 — 0.5Flush skid inlet — 0.25 — 0.25Flush skid outlet — 0.25 — 0.25Total (volume/h)d) 110 l 6.0 2700 l 17.8

a) The calculation is based on the following conditions: feed material conditions: feed concentration,5 g/l; feed volume, 2700 l; feed conductivity, �4ms; feed pH, 7.2.

b) Sartobind1 Q configuration: �4.5m2 [2� 3000 þ 1� 1000 Q membrane (cross-sectional area:3080 cm2)], capacity: 3000 g/m2, linear velocity: 450 cm/h or 23.1 l/min.

c) Q SepharoseTM Fast Flow configuration: 100 cm� 20 cm, 157 l, capacity: 100 g/l. Linear velocity:120 cm/h or 15.7 l/min, residence time: 12.5min.

d) Total volume not including load volume, DV: device volume; CV: column volume.


scale-down model, it has been shown that the axially directed velocity of a mobilesolution ismuch faster in the center of theMAunit than near the edges of adsorptivebed [66]. The decrease in apparent porosity from the center of the bed to the gasketarea or the edges of adsorptive bed may account for the decrease in velocity observedadjacent to the gasket or edges [66,88]. In contrast, radial flow absorbers, prepared byspirally winding a flat sheet membrane over a porous cylindrical core, are used in thecapsules for large-scale processes [70]. Distorted or poor inlet flow distribution hasless effect in a scale-up capsule, thus leading to a smaller operational backpressuredrop for large-scale operation [53,79]. The scalability and operating backpressure,therefore, are main issues for the MA scale-down model. An extremely highoperational backpressure observed with the scale-down model can lead to consider-ably less capacity and an oversized membrane at scale. This over sizing may lead toan erroneous economic calculation. Larger scale membranes do not have back-pressure issues. However, the use of a large MA in a scale-down model is also notpractical for a viral clearance study due to extreme cost of viruses and feedstock. Thetransport phenomena of Q membrane chromatography have been studied[66,70,89–91]. However, most of these reports are based on stacked flat-sheetmembrane units, the so-called scale-down model. The results and predictionsfrom these models might not be very useful to implement MA chromatographyin FT mode.

13.2.2.3 An Improved Scale-down Model and Process CapacityRecently, a new scale-down model, the Q125, was designed for the Q membrane[53,84]. Q125 mimics the liquid flow path found in the large-scale modules thatshows performance parameters comparable to the larger Sartobind1 Q units.With the new scale-down model, a process capacity of >3600 g/m2 (total mem-brane surface) or >13.2 kg/l (membrane volume) at a maximal flow rate of600 cm/h with an operational backpressure of <18 psig was achieved. A viral-log-reduction of >5.6 for MuLV virus was obtained [84]. Since such high processcapacity required a significant amount of feedstock and viral preparations, theviral studies with PRV, MVM, and Reo-3 viruses were not carried out. The viralclearance studies with the four model viruses were run at a flow rate of 450 cm/hand process capacity of 3 kg/m2. The results are summarized in Table 13.10. Thefact that the viral clearance at the high flow rate and high capacity is similar to theMuLV viral log reduction at 450 cm/h and a capacity of 3000 g/m2 (Table 13.10)may indicate that a similar viral clearance can be achieved at higher operationalflow rate and capacity for the other three model viruses. The highest viralclearance capacity so far that we obtained is 6.3 kg/m2 or 23.1 kg/l with aMVM clearance of >6.97 LRV [45].These concepts and parameters have been successfully applied to purify four lots

of recombinant human antibody at the 2000 l scale. The impurities including HCP,DNA, and leachable protein A were below the limit of detection before load on QMA. Figure 13.5 presents a typical viral study result forMVM (1% spike) at 450 cm/hat a 3000 g/m2 capacity with an operational backpressure drop of <16 psi [53,79].The current data is extremely encouraging to consider the replacement of Q-packed


Table 13.10 Viral clearance with Q 125. a), b), c)

Run Loadcondition

MuLV (80–120 nm)

PRV (120–200 nm)

Reo-3 (60–80 nm)

MVM (18–26 nm)

Viralspike

A 3kg/m2 at450 cm/h,�4ms/cm,pH 7.2

�5.35 �5.58 �7.00b) �6.03 1%

B 3 kg/m2 at450 cm/h,�4ms/cm,pH 7.2

�5.52 �5.58 �6.94c) �6.03 1%

C 3.6 kg/m2 at600 cm/h,�4ms/cm,pH 7.2

�5.59 — — — 1%

Viralrecovery(%)

— 70%(n¼ 3)

100%(n¼ 2)

100%(n¼ 2)

100%(n¼ 2)

—

a) Study performed in conjunction with Sartorius Inc., Germany,MuLV:murine leukemia virus; MVM:minute virus of mice; PRV: pseudorabies virus; Reo-3: respiratory enteric orphan III.

b) The study was performed at 8 �C.c) The study was performed at 13 �C.

Load Capacity (l)/125 cm at 450 cm/h at RT

6.17 LogViral Reduction

Maximum Pressure: <15 psi

2

0.0

2.0

4.0

6.0

8.0

10.0

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ress

ure

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)

3

4

5

6

7

8

9

10

11

LRV

MV

M (l

og)

PressureLog Reduction

7.56.04.53.01.50

Figure 13.5 A typical viral study result forMVM (1% spike) at 450 cm/h at a 3000 g/m2

capacity, a maximal pressure of <15 psi and anoperational backpressure drop <16 psi. LRV:log reduction value; MVM: minute virus of

mice; RT: room temperature. Source:Reproduced with permission from BioprocessInternational for the article by Zhou J andTressel T [62].


bed column with MA as a polishing step in a flow-through chromatography modefor process-scale antibody production.

13.2.2.4 Cost AnalysisThus far,Qmembrane inFTmode is theonlydisposable chromatography systemusedin large-scale purification. However, its application in process-scale Mab productionshowsconsiderable impact onpurification [54].Usingcurrent process capacity of 3 kg/m2 formembrane operation, it is apparent that columnoperations aremore expensivethan Q membrane operations [54]. For instance, when WFI price of $0.2 per liter isused for the process, assuming 400 runs of 20 000 l cell culture in a 10-year productionscenario, the cost savings due to use of Q membrane over packed column is about28.7% using newly updated pricing formembrane and resin. Since Qmembrane is adisposable operation system, there are noneeds for columnhardware, packing relatedstudies, resin’s life studies, carryover, assay development, and validation costs [54].It is especially economical when disposable technology is used in the FIH (first in

human or Phase I) process [84]. A cost saving of 62% can be readily achieved sincethe resin is normally only used �10 cycles for Phase 1 and II production.

13.2.2.5 Major Limiting FactorsThe MA devices are not without some issues. Even with improved scale-downdevices, the operational backpressure for Q membrane chromatography alwaysinfluences process capacity. In general, a solution with a high viscosity generates ahigh operational backpressure. However, viscosity of a solution is related to otheroperation parameters, such as operational temperature, pH, and conductivity of thesolution, and protein concentration if the solution contains protein.The viscosity of the buffer and antibody solutions is directly affected by operation

temperature. The effect of temperature on buffer and antibody solution is plotted inFigure 13.6. At a temperature of 4 �C, the viscosity of the antibody solution (3mg/ml

Temperature (°C)

0.8

0.9

1.0

1.1

1.2

1.3

1.4

1.5

Vis

cosi

ty (c

P)

5

10

15

20

25

30

Bac

kpre

ssur

e (p

si)

Protein Viscosity (cP)Buffer Viscosity (cP)Backpressure (psi)

4 8 13 20Temperature (°C)

0.8

0.9

1.0

1.1

1.2

1.3

1.4

1.5

Vis

cosi

ty (c

P)

5

10

15

20

25

30

Bac

kpre

ssur

e (p

si)

Protein Viscosity (cP)Buffer Viscosity (cP)Backpressure (psi)

4 8 13 20

Figure 13.6 Effects of temperature onviscosities of buffer and antibody solution(3mg at pH 7.2). Vis: viscosity. Back pressure

obtained with Sartobind1 Q-125. Buffer testedwas 20mM Tris contained 25mM NaCl, pH 7.2[49].


at pH 7.2 and about 4ms/cm) was found to be much higher than the viscosity at8 and 13 �C [45,84]. The viscosity of an antibody solution was found to be muchhigher than the viscosity of the control buffer solution at the same low temperature.High viscosity at low temperature significantly increases operational backpressureand results in lower flow rates. Thus, temperature is the most important limitingfactor in MA operation.While anantibodysolutionat low temperature loadedontoQmembranegenerates a

high backpressure (Figure 13.7), it has no effect on viral clearance (Table 13.11) [84]. Ata low temperature, even smaller particles can generate significant backpressure.Figure 13.8 illustrates insignificant backpressure changes when four model viruseswith different particle sizeswere processed for a viral study at pH7.2, 4ms/cm, 20 �C,and a process capacity of 3000 g/m2. Indeed, at room temperature, the operationalbackpressure is independent of virus particle sizes (18–200nm, a similar back-pressure curve is found for Reo-3, PRV, and MuLV viral particles). Figure 13.9 plotsa linear relationship between operational temperature and backpressure, indicatingthat tomaintain the least operational backpressure, theQmembrane chromatographyshould be operated at 20 �C or above [84].

Process Volume (l)

0

5

10

15

20

25

30

35

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 2 4 6 8 10

MVM 20 nm RTMVM 20 nm 8oCMVM 20 nm 13oC

Process Volume (l)

0

5

10

15

20

25

30

35

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 2 4 6 8 10

MVM 20 nm RTMVM 20 nm 8oCMVM 20 nm 13oC

Figure 13.7 Temperature effect onbackpressure with Q membrane (Sartobind1

Q125). Viral study with MVM viruses

(18–20 nm sizes) at pH 7.2, �4ms/cm with aprocess capacity of 3000 g/m2. MVM: minutevirus of mice; RT: room temperature.

Table 13.11 Backpressure by low temperature affects the capacity but viral clearance.a)

Run MMV (LRV) Operational temperature (�C) Pressure (psi)

1 6.03 8 292 6.03 13 223 6.12 20 13

a) Process capacity: 3 kg/m2.


The second limiting factor is found when the buffer pH is close to the pI value ofprotein. Figure 13.10 indicates that the viscosity appears to be maximal at the pI ofthe molecule studied [84]. Other limiting factors including conductivity need to bemonitored. A low conductivity results in a high molecular interaction between theantibodies, thereby increasing solution viscosity.

13.2.2.6 Robust Viral RemovalPublished data, using the Q membrane with different feed materials, demonstratesthat Q membrane chromatography is a robust viral removal step [54,84,87].

0

5

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15

20

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 2 4 6 8 10Process Volume (l)

Reo-3 80 nm RTPRV 180 nm RTReo-3 80 nm RTPRV 180 nm RTMuLV 100 nm RTMVM 20 nm RT

0

5

10

15

20

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 2 4 6 8 10Process Volume (l)

Reo-3 80 nm RTPRV 180 nm RTReo-3 80 nm RTPRV 180 nm RTMuLV 100 nm RTMVM 20 nm RT

Figure 13.8 Particle size effect onbackpressure with Q membrane (Sartobind1

Q125). Viral study performed at �20 �C for awith four model viruses of different particlesizes at pH 7.2, �4mS/cm with a process

capacity of 3000 g/m2. MuLV: murine leukemiavirus; MVM: minute virus of mice; PRV:pseudorabies virus; Reo-3: respiratory entericorphan III; RT: room temperature.

0

5

10

25

35

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 5 10 15 20 25Temperature (°C)

15

20

30

y=-1.4287x + 40.151R2=0.9908

0

5

10

25

35

Ope

ratio

nal B

ackp

ress

ure

(psi

)

0 5 10 15 20 25Temperature (°C)

15

20

30

y=-1.4287x + 40.151R2=0.9908

Figure 13.9 A linear relationship exists between operational temperature and backpressure.


We found that no adverse effects on the viral removal capacity of themembranes wasfound with flow rates from 240 to 600 cm/h. Consistent viral removal results wereachieved with process capacities ranging from 49 to 6310 g/m2 after spike with thenonenveloped DNA model virus, MMV [45].Mass balance data for a viral clearance study is extremely important to demon-

strate an efficient virus removal by a MA device (Table 13.10). For example, 100%recovery was obtained for PRV (n¼ 2), Reo-3 (n¼ 2), and MVM (n¼ 2) virusparticles with different particle sizes from 18 to 200 nm when the membranewas stripped with 1M NaCl, demonstrating efficient charge capture for the threemodel viruses. When theMAwas treated with high salt, an average of only 70% viralrecovery was obtained for MuLV virus [n¼ 3].

13.2.2.7 Lot-to-Lot VariabilityA lot-to-lot variability studies are important to ensure consistent operation. Recentdata using X-MuLV and MMV indicate that the viral clearance performance isreliable and consistent with different lots of the membrane devices [45].

13.2.2.8 Concern for Qualification and Membrane ReuseA sensitive detection of any defects in the Qmembrane device could still be an issue,particularly, when the device is used at a high flow rate with a short bed height in aflow-throughmode. Regular asymmetry (As) and height equivalent to the theoreticalplate (HETP) qualification can examine the separation power; however, it isextremely difficult to detect microdefects in flow-through mode. The same conceptapplies to column operation: As andHETP cannot detect microchanneling in a flow-through mode of a defined column. However, it is safer when a column is operatedat relative slow flow rate since its bed height is much larger than a Q membranedevice’s. Although the current pressure or bubble testing can be used for Q

50

60

70

80

90

100

110

120S

EC

Pur

ity (%

)

0.84

0.86

0.88

0.90

0.92

0.94

0.96

0.98

Vis

cosi

ty (c

P)

SEC PurityViscosity

pH 7.05 pH 7.20 pH 7.45 pH 7.85 pH 8.0550

60

70

80

90

100

110

120S

EC

Pur

ity (%

)

0.84

0.86

0.88

0.90

0.92

0.94

0.96

0.98

Vis

cosi

ty (c

P)

SEC PurityViscosity

pH 7.05 pH 7.20 pH 7.45 pH 7.85 pH 8.05

Figure 13.10 Antibody’s dimerization andsolution viscosity at solution’s pH near the pI ofthe antibody. When the buffer pH is close to thepI value of protein, the antibody in this buffer

does not generate dimer or aggregates butgenerates a higher viscosity (this antibody has apractical pI of 7.5) [49].


membrane integrity qualification and model proteins can be used for bindingcapacity determination, a more sensitive detection for defects, particularly, for futurelarge-scale manufacturing may be desired [45].To usemembrane chromatography devicemore economically, membrane reuse is

proposed. Our major concerns are (i) inconsistent operational pressure from one toanother run even the same membrane device is used; (ii) sensitive qualification hasto fit the purpose of membrane reuse; and (iii) adequate validation includingresidual protein and virus carryover, cleaning, and storage validation.Membrane reuse currently is attractive mainly because of the pricing of the

device, minimal usage of buffers and tanks, and high product throughput.Although we observed significant variability of operational pressure betweenruns, solution to it using proper cleaning and storage seems achievable. Tomeet the purpose of membrane reuse, the sensitive qualification method hasto be user-friendly. Physical detection such as using helium gas seems not feasiblebecause of its complicated settings and need of expensive apparatus. Therefore,sensitive detection using chemicals is perhaps the attractive choice. Validation isthe major concern, and the major benefits of disposable system are left out whenthe membrane is reused.

13.2.3Implementation of Other Membrane Technology in AntibodyLarge-Scale Purification

Besides Q membrane development, other membrane devices also gained remark-able progress in the past decade [63,64,66,92–95]. Currently, affinity, hydrophobicinteraction (HIC), and CEX-based membrane chromatography in bind/elutionmode is in development [96]. According to their possible operation modes, theadvantages and disadvantages for each membrane chromatography in Mab purifi-cation summarized in Table 13.12. Since operation of ion exchange membrane orresin-packed column is residence time independent, a fast operational flow rate isachievable. Although AEX membrane in flow through mode and CEX in bind/elution mode can be used for similar purpose, such removal of DNA, viruses,endotoxin, and some host cell proteins, it is clear that a significant improvement isrequired for CEX in bind/elution mode due to membrane’s poor binding capacity. Apoor operational output is expected for AEX in bind/elution mode even though thedevice can be utilized for Mab with a low pI such in the case of IgG4 purification.CEX membrane chromatography in flow through mode should be useful forremoving aggregates and dimers when the loading condition is carefully defined.In this case, the advantage of high flow rate is readily applied in the process.While a high flow rate can not be used for resin-time-dependent operation unit

such as HIC and affinity membrane devices, removal of Mab aggregates and dimerscan be expected using HIC in a flow through mode. A fast analytical titerdetermination could be valuable with an affinity membrane device in bind/elutionmode. Cost might be a concern when the affinity membrane is used for either bind/elution or flow through mode.


13.2.4Future Perspectives

The antibody production technology advances rapidly. Cell culture titer continues toincrease in a significant way, and purification process capacity and processing timecontinue to improve. Innovation is the engine leading to continual enhancement inupstream and downstream process optimization.New disposable systems in upstream and downstream processing are widely

accepted as an economical choice. New high capacity and high-throughput equip-ment and devices are entering process development toolboxes at a fast pace. It isexpected that COGs will be reduced dramatically at all levels of process developmentand at large-scale antibody production in the coming years.

Acknowledgment

The authors sincerely acknowledge the following companies and teams: Sartorius,Inc., Millipore, Inc., Pall Life Sciences, and Asahi Kasei Corporation, Amgenteams in Thousand Oaks: Purification Process Development, Pilot Plant Opera-tions, Cell Science & Technology, and Analytical & Formulation Sciences as well asthe Amgen Bio-safety Laboratory in Seattle. The authors also wish to thank DrSam Guhan for his strong support in the area of technology evaluation anddevelopment.

Table 13.12 Consideration of membrane technology in antibody purification.

Type Opmode

BCap

Res. time Function Feasibility

AEX FT Excl Independent Removes DNA, viruses, endotoxin,HCP

Excl

AEX B/E Poor Independent Mab with low pI, removes endotoxin,and HCP by selectivity

Poor

CEX FT Excl Independent Removes aggregate and dimers,some HCP

Good

CEX B/E Poor Independent Removes DNA, viruses, endotoxin,HCP

IR

HIC FT Excl Dependent Removes aggregate and dimers,some HCP

OK

HIC B/E Poor Dependent Removes aggregate and dimers,some HCP

Poor

Affinity FT Excl Dependent Removes specific antigens Poor/costAffinity B/E Poor Dependent Removes most impurities, DNA,

virusOK/anal.

Excl, excellent; IR, improvement required; anal, analytical usage.


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