The market for therapeutic monoclonal antibodies(mAbs) has grown tremendously in the last decade,and it is estimated that mAbs and their derivatives

account for almost 36% of the biopharmaceuticals underdevelopment, including vaccines and gene therapy (1).These and other therapeutic proteins are produced at theindustrial scale using various recombinant cell lines, suchas bacteria (e.g., E. coli), yeast, and mammalian cells (e.g.,Chinese Hamster Ovary (CHO) cells (2)), with CHO cellsbeing the most popular choice because they offer severalbenefits (3).

Biopharmaceutical products must have very high puri-ty, with the concentration of host cell proteins and DNAreduced to the range of parts per million relative to thedesired product, or lower. The final product must also besterile (no viable micro-organisms present), and shouldcontain less than 10 ng of DNA per dose and less than onevirus per million doses. This stringent purification ofmAbs produced in CHO cells is typically accomplishedusing a three-column chromatography process (Figure 1)that consists of protein A affinity chromatography as aninitial capture step, followed by cation exchange (CEX)and anion exchange (AEX) chromatography as polishingsteps and a virus filtration (VF) step (3). The protein Aligand has a high affinity for one area of the mAb, specifically the crystallizable fragment (Fc), which enablesthe mAb’s capture from the cell culture fluid (thus theterm affinity chromatography).

Even though these chromatography steps are able tomeet the stringent purification requirements, they areexpensive, especially the protein A affinity step, whichaccounts for almost 35% of the total raw material costs for

downstream purification (4). With growing demand fortherapeutic mAbs and increasing market competition, sig-nificant attention is being focused on reducing manufactur-ing costs and improving process efficiency for industrial-scale production. As a result, there has been much interestin the development of cost-effective non-affinity purifica-tion processes that do not involve protein A.

In one of the early studies on the purification of indus-trial-scale mAb feed streams using a non-affinity process(4), Fahrner and Follman integrated different chroma-tography steps into three-column non-affinity purificationsequences and evaluated their performance. Three differentnon-affinity sequences were identified that provided prod-uct quality comparable to the affinity process. This three-column non-affinity process with ultrafiltration/diafiltration(UF/DF) has since been streamlined into a two-columnnon-affinity process with the integration of high-perform-ance tangential-flow filtration (HPTFF) (5). This two-column non-affinity process utilizes cation exchange as acapture step, followed by anion exchange as a polishingstep, followed by HPTFF (which allows concentration, formulation and purification in a single step). This elimi-nates the most expensive chromatography step (protein Achromatography).

In addition to reducing raw material costs, the two-col-umn non-affinity process can provide several other bene-fits, including a smaller plant footprint and the need forfewer pieces of process equipment (e.g., tanks, columnsand control systems) due to the smaller number of chro-matography steps.

This article addresses several aspects of the two-col-umn non-affinity process: its development and optimiza-

Eliminating the affinity chromatography step in the mAb purification process trims equipment requirements,

shrinks the plant footprint — and reduces costs.

Amit Mehta, Martha Lovato Tse, Jace Fogle, Amy Len, Roshan Shrestha, Nuno Fontes, Bénédicte Lebreton, Bradley Wolk, Robert van Reis

Genentech

PURIFYING THERAPEUTICMONOCLONAL ANTIBODIES

S14 SBE Special Section ❙ Bioprocessing CEP

tion using a manufacturing-scale feed stream; processyield and product quality; virus removal; and applicationto multiple mAb feed streams.

Purification process development Industrial-scale production of therapeutic mAbs is com-

plex, involving process development, process characteriza-tion and validation, process scale-up, automation, qualityassurance, and regulatory compliance. The development ofpurification processes (typically at the laboratory scale)involves optimization and integration of several unit oper-ations that provide the desired purity, product quality,throughput and yield.

Although a comprehensive discussion of the differentunit operations associated with the two-column non-affini-ty process is beyond the scope of this article, the keyaspects related to the development and optimization ofeach step are addressed in the following sections. Processdevelopment is illustrated using, as an example, the pro-posed two-column non-affinity process to purify a mAb Afeed stream. mAb A has an iso-electric point (pI) of 9.3and molecular weight of approximately 145 kDa. HPTFFexperiments were performed on a scale-down control sys-tem using 0.1-m2 prototype membranes with 100–300-kDamolecular-weight cut-off pore size.

CEX column chromatography Cation exchange resins have negatively charged func-

tional groups immobilized on their surface. They can beused as an initial mAb capture stepby loading harvested cell culturefluid (HCCF) under conditionswhere the mAb is positivelycharged (the pH of the materialbeing loaded is lower than the iso-electric point of the mAb). Most ofthe impurities flow through thecolumn during the loading phase;other impurities are separated fromthe mAb during the wash and theelution phases. During the washphase, the column is rinsed with abuffer having an appropriate pHand ionic strength, which removesresidual impurities but leaves themAb bound to the column. A lin-ear salt gradient is typically usedduring the elution phase, whichallows separation of mAb andimpurities, such as aggregates,

■ Figure 1. A three-column affinity process with ultrafiltration/diafiltration (UF/DF) is used topurify monoclonal antibodies expressed from CHO cell lines.

Harvest Protein A CIEX VF AIEX UF/DF

This steputilizes

centrifuge anddepth filters

for removal of intact cells

and cell debris from cell

culture fluid

Removes>98% ofCHOP,

virus, DNAand smallmolecules

Removesproductvariants,CHOP,

DNA, virus,leachedprotein Aand smallmolecules

> 4 logvirus

removal

Removestrace

levels ofCHOP,

DNA andpotentialviruses

Used forproduct

concentrationand

formulation

Low-pH Virus Inactivation

CEP SBE Special Section ❙ Bioprocessing S15

GLOSSARY

Chinese Hamster Ovary proteins (CHOP) — host cell proteinsproduced by CHO cells during fermentation, in addition to thedesired protein product

Diavolume — the ratio of total volume of buffer added duringthe operation divided by the retentate volume

Diafiltration (DF) — a technique that utilizes a semi-permeablemembrane to exchange the product of interest from one liquidmedium into another

Dynamic Binding Capacity (DBC) — amount of target proteinthat binds to the media under actual flow conditions prior toproduct breakthrough

Harvested Cell Culture Fluid (HCCF) — cell culture fluid containing product of interest and devoid of intact cells and cell debris

High-Performance Tangential-Flow Filtration (HPTFF) — anemerging technology that separates proteins by exploiting thedifferences in size as well as charge

Iso-electric Point (pI) — pH at which there is no net charge onthe protein

Load Density — amount of product loaded per volume of media(resin, membrane)

Monoclonal Antibody (mAb) — very specific, identical antibodymolecules that are produced by a cellular clone

Process Throughput — amount of mAb processed per unit of time

Ultrafiltration (UF) — a technique that utilizes a semi-permeable membrane to concentrate the product of interest

product variants and CHO protein (CHOP), due to theirdiffering electrostatic affinities for the charged resin. Thus, the development of a CEX step is very rigorous andrequires a thorough study of several process parameters. It involves: the selection of appropriate chromatographicresin; the optimization of load density, load flowrate (residence time at constant bed height), load pH, and ionic strength; and the optimization of the wash and elution phases.

The selection of an optimal CEX resin entails thestudy of the inter-relationships between load pH, ionicstrength, load density (mass of mAb loaded divided byresin volume), impurity removal, load flowrate, and pro-cessing time required for a given feed stream. For example, loading the HCCF at lower pH and ionicstrength may provide higher dynamic binding capacity(DBC) (6), but may significantly increase the processingtime if a large dilution volume is required to adjust thepH and ionic strength. Lowering the load density mightenhance impurity removal, but would increase costs dueto lower utilization of chromatography resin. In additionto process parameters, the performance of CEX resin canbe strongly affected by product characteristics, so differ-ent resins would be optimal for different products. Anideal CEX resin should provide sufficient DBC (typically> 50 g/L), high yield (> 90%) and acceptable productpurity in the eluent (i.e., low levels of impurities such ashost cell proteins, DNA, aggregates, etc.) to ensure prod-uct stability and reasonable cycle time.

To identify the optimal CEX resin for the mAb A feedstream, several resins were screened. The DBC of eachresin was evaluated by loading the HCCF while varying

the pH, conductivity and flowrates. Figure 2 shows theDBC as a function of load conductivity for different CEXresins at pH 5. The DBC varied from 2 to 129 g/L at 1%product breakthrough.

The DBC increased with decreasing load conductivitydue to enhanced electrostatic interaction between the posi-tively charged mAb and the negatively charged resin. Thebinding capacity also varied significantly from one resinto another at any given pH and conductivity. This can beattributed to differences in resin characteristics, such aspore size, ligand density, functional group and spacer arm.

A maximum DBC for Resin A was observed at approximately 8 mS/cm, and can be attributed to thereduction in protein uptake by the CEX resin due tocharge exclusion (6). The DBC maximum depends onmedia characteristics, such as charge density and poresize, and solution conditions, such as pH and ionicstrength (6). It is likely that a similar trend would beobserved with other resins if a wider conductivity range(i.e., less than 6 mS/cm) were explored.

The wash and elution phases were designed by loadingthe resins to 80% of their DBC and evaluating the effectof parameters such as pH, ionic strength and bufferspecies on impurity removal. CHOP levels in the elutionpools were determined using an enzyme-linkedimmunosorbent assay (ELISA), while mAb concentrationwas determined by measuring absorbance at 280 nm on aUV/visible spectrophotometer. Table 1 summarizes theCEX step yield and CHOP levels in the elution pools atthe optimal wash and elution conditions for each resin.Most of the resins provided more than 95% product yield,but CHOP clearance varied significantly from one resin to another.

The best resin was selected based on binding capacity,pool purity, load flowrate (processing time) and the abilityof the downstream unit operations to remove residual impu-rities in the CEX elution pool. For the mAb A feed stream,resin C was chosen, because it provided significantly higherimpurity removal than the other resins with adequate bind-ing capacity and desired downstream purification.

■ Figure 2. Dynamic binding capacity for cation exchange resinsused to purify the mAb A feed stream as a function of load conductivity at pH 5.

Resin AResin BResin C

Resin EResin D

4 6 8 10 12 14 16Conductivity, mS/cm

Bin

ding

Cap

acity

, g/L

0

20

40

60

80

100

120

140Table 1. Yield and CHOP levels in the elution pool at the

optimal wash and elution conditions for CEX resins. Resins were loaded to 80% of their DBC. CHOP levels in

the initial load were approximately 140,000 ng/mg of mAb.

CEX Resin Load g/L Yield, % CHOP, ng/mg

Resin A 77 97 6,600Resin B 81 93 10,000Resin C 69 98 2,600Resin D 87 98 6,800Resin E 41 — 5,000

S16 SBE Special Section ❙ Bioprocessing CEP

AEX chromatography Anion exchange resin chromatography is typically used

as a second or third chromatography step in traditionalmAb purification processes to remove trace levels ofCHOP, nucleic acids, endotoxins and any viruses that maybe present. This step typically operates in a flow-throughmode, that is, under conditions where the mAb is positive-ly charged and flows through the column while the nega-tively charged impurities bind to the positively chargedresin. Although AEX resins are a powerful purificationtool, their main drawback is a low mass-transfer rate dueto intra-pore diffusion, which results in lower processthroughput (mass of mAb processed per unit of time). As aresult, in some cases, the process-scale columns need to beoversized simply to meet the throughput requirements.

Membrane chromatography is an attractive alternativeto conventional resins, because binding sites are locatedwithin the convective path of the fluid, which enableshigher throughput (7). In addition, the availability ofmembrane adsorbers in pre-packed formats significantlysimplifies manufacturing by eliminating the need for pack-ing and unpacking columns. Because of these key bene-fits, AEX membrane chromatography was selected as thesecond step for the non-affinity process. Because it oper-ates in a flow-through mode, process development forAEX chromatography is less intensive compared to CEX.It involves the selection of the optimal media, and thedetermination of process parameters such as load pH andconductivity, flowrate, and load density.

AEX membrane screening was done by loading theCEX elution pool under different solution conditions andanalyzing the flow-through stream for impurities, such as

CHOP and DNA. The load pH was kept at 1–1.5 pH unitbelow the mAb iso-electric point, as the mAb maintains aweak positive charge under these conditions and flowsthrough the membrane. Low load conductivity is usuallydesired, as CHOP removal typically increases withdecreasing conductivity due to enhanced electrostaticinteraction between the CHOP and the membranes.However, lower load conductivity involves larger dilutionof the load, resulting in larger load volume and longer pro-cessing time. (The effect of conductivity on processingtime and impurity removal was studied in detail prior tothe selection of the process parameters.)

In addition, the pressure drop across the membraneswas monitored, because membrane adsorbers can be oper-ated safely only within a certain pressure range (specifiedby the vendor). Figure 3 shows the differential pressuredrop across two commercially available membranes as afunction of the load density. Unlike Membrane B, thepressure drop across Membrane A increased dramaticallywith increasing load density. If the safety limit for thepressure drop is 40 psi, the data suggest that Membrane Acan only be loaded to 1 kg/L, while Membrane B can beloaded to > 5 kg/L. Membrane adsorber B was selectedfor purification of the mAb A feed stream because of itsbetter pressure performance. Membrane B also providedapproximately a twenty-fold CHOP reduction at a loaddensity of 10 kg/L.

High-performance tangential-flow filtrationHPTFF (Figure 4) is an emerging membrane tech-

nology that performs protein purification, concentrationand formulation in a single step. It utilizes a positively

■ Figure 4. HPTFF is a membrane process used to separate mAband CHO proteins.

++

++

++

++

+

+ +

+ +

+

++

++

RetentateFlow

Filtrate(To Drain)

PositivelyCharged

Membrane

CHOP

mAb

■ Figure 3. Pressure drop across two commercially availablemembranes used to purify the mAb A feed stream as a function of load density.

Membrane A

Membrane B

0 1 2 3 4 5Load Density, kg/L

0

20

40

60

80

100

120

Pre

ssur

e, p

si

Conditions:pH = 8Conductivity = 4 mS/cm

CEP SBE Special Section ❙ Bioprocessing S17

charged, large-molecular-weight cut-off (~100–300 kDa)membrane (whereas traditional UF/DF uses a 10–30-kDaneutral membrane). The proteins are separated by exploit-ing differences in their size as well as charge (5–8).HPTFF performance is governed by membrane properties,such as pore size and membrane surface-charge density,and process parameters, such as filtrate flowrate, numberof diavolumes, load pH and ionic strength.

Experiments were performed with a prototype mem-brane having optimal pore size and charge density todetermine process conditions that would provide a yieldgreater than 95% for this step and CHOP levels less than10 ng of CHOP per mg of mAb in the final bulk. Solutionconditions were such that the positively charged mAb wasretained by the membrane while neutral or weakly charged

CHOP passed through it. Figure 5 shows the observedsieving coefficients (So) of mAb A and CHOP as a func-tion of the number of diavolumes at pH 5.3 and conductiv-ity of 4 mS/cm. So is the ratio of protein concentration inthe filtrate to that in the retentate, and serves as a measureof protein transmission through the membrane. Filtrateflux was set at 75 L/m2-h during ultrafiltration and at 50L/m2-h during diafiltration. The observed sieving coeffi-cient of mAb (~ 0.001) was approximately two orders ofmagnitude lower than that of the CHOP during the entireultrafiltration and diafiltration phase, enabling better than95% product yield to be achieved after 20 diavolumes.The effect of the mAb sieving coefficient on yield for adiafiltration step with 20 diavolumes is shown is Figure 6.The product yield decreases dramatically as the sievingcoefficient increases. For example, the product yielddecreases from approximately 90% at So = 0.005 to lessthan 15% at So = 0.1.

Figure 7 shows the CHOP levels in the retentate duringultrafiltration and diafiltration. While the mAb is electro-statically excluded from the membrane pores, CHOPpasses through, so CHOP levels decrease with increasingdiavolumes. CHOP concentration in the initial load wasapproximately 120 ng/mg, and was reduced to 30 ng/mgduring ultrafiltration. CHOP levels were further reduced to10 ng/mg after 4 diavolumes and to less than 5 ng/mgafter 10 diavolumes.

Virus clearanceCHO cell culture fluid typically contains 103 to 109

retrovirus-like particles per mL, and since the industrystandard is less than one retrovirus-like particle per million doses, downstream purification needs to provideseveral orders of magnitude of retrovirus removal andinactivation. Conventional purification processes typically

■ Figure 7. CHOP levels in the retentate as a function of the number of diavolumes during diafiltration of the mAb A feed stream.

Number of Diavolumes

0 5 10 15 200

20

40

60

80

100

120

140

CH

OP,

ng/

mg

UF DF

Conditions:mAb A concentration in the bulk fluid = 40 g/LFiltrate flux = 50 L/m2-hLoad pH = 5.3Conductivity = 4 mS/cm

0.001 0.01 0.1Observed Sieving Coefficient, So,mAb

100

80

60

40

20

0

Yie

ld, %

■ Figure 6. Effect of the mAb sieving coefficient on yield [Y = exp(–N×So) (9), where N is the number of diavolumes and So is the observed sieving coefficient for the mAb] for a diafiltration step with 20 diavolumes.

■ Figure 5. Observed sieving coefficients for mAb A and CHOP asa function of the number of diavolumes during diafiltration of themAb A feed stream.

Obs

erve

d S

ievi

ng C

oeffi

cien

t, S

o 1

0.1

0.01

0.001

0.00010 2 4 6 8 10 12 14 16 18 20

Number of Diavolumes

mAb ACHOP

Conditions:mAb A concentration in the bulk fluid = 40 g/LFiltrate flux = 50 L/m2-hLoad pH = 5.3Conductivity = 4 mS/cm

S18 SBE Special Section ❙ Bioprocessing CEP

target a 12–18 log reduction value (LRV) for retrovirusremoval (i.e., the level of retroviruses is reduced by 12–18orders of magnitude).

Several technologies, such as chromatography, virus filtration, low-pH inactivation and heat inactivation, areused to meet the viral safety requirements. Table 2 lists the degree of retrovirus removal and inactivation achievedby the various steps of the affinity and the two-columnnon-affinity processes.

The use of cation exchange as the initial capture stepand the incorporation of a viral heat inactivation step inthe two-column non-affinity process compensate for theloss of retrovirus removal that would have been achievedby the protein A chromatography and the low-pH inactiva-tion steps that were eliminated. Analytical tests were performed to ensure that viral heat inactivation had noadverse impact on product quality. Both the affinity and

the non-affinity processes provide approximately 18 logsof retrovirus removal/inactivation.

In the affinity process, removal of adventitious virus isachieved in the protein A and anion exchange chromatog-raphy steps. Adventitious-virus removal in the non-affini-ty process can be obtained by utilizing new technologies,such as parvovirus filtration and UV inactivation, alongwith traditional anion exchange chromatography.

Process yield and purity After the individual unit operations were developed

and optimized at the laboratory scale, the two-chromatog-raphy-step non-affinity process was tested at the pilot scaleusing the mAb A feed stream. The process yield and theCHOP removal data across different steps are summarizedin Table 3. Process yield was approximately 80%, andCHOP levels in the final pool were less than 5 ng/mg forall of the experimental runs. Both process yield and CHOPlevels in the final pool are comparable to those obtainedwith the conventional affinity process. Other impurities,such as DNA, aggregates and small molecules, also metpurity specifications for therapeutic mAbs.

The non-affinity process with other mAbsThe data presented so far illustrate the application of

the two-column non-affinity process to a single mAb feedstream. In order to study the broad application of this

process, different mAb feed streamswere obtained and purified.Although the overall purificationsequence (CEX as a capture step,AEX as a polishing step, andHPTFF for concentration, formula-tion and purification) remained thesame from one mAb to another,individual media and processparameters were tailored for different mAbs.

Table 4 summarizes the CHOPremoval across different steps withthree other mAb feed streams. The

Table 2. Removal and inactivation of retroviruses by different unit operations

in the affinity and non-affinity processes.

Affinity Process Non-Affinity ProcessVirus Virus

Clearance Clearance Unit Operation (LRV)* Unit Operation (LRV)*

Protein A 2.0 Cation Exchange† 3.8Low-pH Inactivation 6.0 Heat Inactivation† ≥4.0Virus Filtration 4.6 Virus Filtration 4.6Anion Exchange 5.5 Anion Exchange 5.5

Process LRV 18.1 Process LRV ≥17.9

* LRV = log reduction value = orders of magnitude reductionachieved† LRV across the CEX and the viral heat inactivation steps was obtained using mAb A feed stream. LRV across the other steps is based on internal R&D data or data supplied by vendors.

Table 4. Amount of CHOP removed (ng/mg) by different steps in the non-affinity process

with different mAb feed streams.

Step mAb B mAb C mAb D

HCCF 1,470,000 240,000 500,000CEX 145,000 530 450AEX 410 15 19HPTFF 2 < 1 2

Table 3. Process yield and CHOP clearance data (ng/mg) for different steps of the two-column non-affinity process.

Step Run 1 Run 2 Run 3 Run 4 Run 5 Run 6 Run 7

InitialLoad 162,000 104,000 104,000 140,000 76,000 76,000 67,000

CEX 3,000 1,400 1,300 3,400 1,400 1,150 950AEX 90 180 120 230 150 120 70HPTFF 4 4 5 3 4 3 3

Overall ProcessYield 82% 79% 79% 76% 78% 81% 82%

CEP SBE Special Section ❙ Bioprocessing S19

CHOP concentration in the final pool for all mAbs was 2ng/mg or less. Even though the CHOP level in the mAb BHCCF was almost fivefold to sevenfold higher than those inother feed streams, the CHOP concentration in the HPTFFpool was, nevertheless, reduced to 2 ng/mg, illustrating therobust purification capability of the non-affinity process.

Concluding remarksThe two-chromatography-step non-affinity process

that uses cation exchange as a capture step and HPTFF as an additional purification tool shows much promise for the purification of therapeutic monoclonal antibodies.This process reduces the number of processing steps aswell as total costs by eliminating the most expensive stepin the current antibody manufacturing process — proteinA affinity chromatography. Yield, purity and product quality comparable to the traditional affinity process were obtained at pilot scale using a commercial-scale feed stream. The non-affinity process was successfullyemployed for multiple feed streams, illustrating its potential as a future platform process for mAb purifica-tion. Furthermore, this process uses technologies that canbe retrofitted to existing manufacturing plants, eliminatingcapital expenditure for high-volume production.

AMIT MEHTA is a group leader and an engineer in the Late Stage PurificationDept. at Genentech, Inc. (South San Francisco, CA; Phone: (650) 225-1087;E-mail: mehta.amit@gene.com). His research interests include membraneseparations and biopharmaceutical process development. He holds a PhDin chemical engineering from the Pennsylvania State Univ.

MARTHA LOVATO TSE is a scientist in the Early Stage Purification Dept. atGenentech, where her work focuses on developing purification processesfor molecules in early stages of development. She received a PhD inchemistry from The Scripps Research Institute under the direction of Prof.Paul Schimmel and was a National Science Foundation Postdoctoral Fellowat Stanford Univ. under the direction of Prof. Chaitan Khosla.

JACE FOGLE is an engineer in the Process Research and Development Dept. atGenentech. He received a BS in chemical engineering from thePennsylvania State Univ. and an MS and a PhD from the Univ. of Virginia.

AMY LEN is an engineer in the Late Stage Purification Dept. at Genentech. Shereceived BS and MS degrees in chemical engineering from Stanford Univ.and is a member of AIChE, ACS and the Society of Women Engineers (SWE).

ROSHAN SHRESTHA is an associate engineer in the Manufacturing Scienceand Technology Dept. at Genentech. He received a BS in chemicalengineering from Stanford Univ.

NUNO FONTES is a group leader and senior engineer in the Late StagePurification Dept. at Genentech. His research interests include proteinpurification and biopharmaceutical process development. He received aPhD from the New University of Lisbon in Portugal.

BÉNÉDICTE LEBRETON is a senior group leader and scientist in the EarlyStage Purification Dept. at Genentech. She received a PhD in chemicalengineering from the Univ. of Birmingham, U.K., is a member of AIChE andACS, and has coauthored 12 publications.

BRADLEY WOLK is a distinguished engineer and director of processdevelopment engineering at Genentech. His research interests includeequipment design and the evaluation of new technologies in the areas ofcell culture, purification and filling. He received a BS in chemicalengineering from the Univ. of California at Davis, and is a member of AIChEand NACE International.

ROBERT VAN REIS is a distinguished engineer and senior group leader in theLate Stage Purification Dept. at Genentech. His research interests includethe development of protein purification technologies to enable cost-effective industrial-scale implementation of manufacturing processes thatproduce pharmaceuticals to meet critical unmet medical needs. Hereceived his education in chemical engineering at the Royal Institute ofTechnology in Stockholm, Sweden, and conducted his thesis research atthe American Red Cross Biomedical Engineering Laboratory in Bethesda,MD. He holds 14 patents and has coauthored 20 journal articles and sevenbook chapters. He has given lectures at AIChE, ACS and North AmericanMembrane Society (NAMS) meetings, as well as at Rensselaer PolytechnicInstitute, Stanford Univ., Univ. of Wisconsin, and Washington State Univ.

CEP

Acknowledgements

We thank Genetech’s Analytical Operations for performing the CHOP assays,and Bin Yang and Min Zhang in the Virus Testing and Development group forproviding the virus removal data. We also thank Arick Brown, Jerry Bill, AiLin and Rachel Dinges in Late Stage Purification for assistance on differentunit operations during pilot-scale runs. And we are grateful to the scale-down and pilot plant groups for supporting the runs. Finally, we acknowl-edge Millipore, Inc. (Bedford, MA) for providing the HPTFF membranes.

Literature Cited

1. Sinclair, A., “A Practical Guide to BiopharmaceuticalManufacturing,” Scrip Reports, London, U.K., and InformaHealth Care USA, New York, NY (2006).

2. “SBE Special Supplement: From Chinese Hamsters toTherapeutic Proteins,” Chem. Eng. Progress, 103 (10), pp. 33–52 (Oct. 2007).

3. Fahrner, R. L., et al., “Industrial Purification ofPharmaceutical Antibodies: Development, Operation andValidation of Chromatography Processes,” Biotech. andGenet. Eng. Rev., 18, pp. 301–327 (2001).

4. Follman, D. K., and R. L. Fahrner, “Factorial Screeningof Antibody Purification Processes Using ThreeChromatography Steps without Protein,” J. Chromatogr. A.,1024, pp. 79–85 (2004).

5. van Reis, R., “Charged Filtration Membranes and UsesTherefor,” U.S. Patent No. 7,001,550 (2006).

6. Harinarayan, C., et al., “An Exclusion Mechanism in Ion Exchange Chromatography,” Biotech. Bioeng., 95,pp. 775–787 (2007).

7. Zhou, J., and T. Tressel, “Basic Concepts in Q MembraneChromatography for Large-Scale Antibody Production,”Biotechnol. Prog., 22, p. 341 (2006).

8. van Reis, R., et al., “High-Performance Tangential FlowFiltration Using Charged Membranes,” Journal ofMembrane Science, 159, pp. 133–142 (1999).

9. van Reis, R., and S. Saksena, “Optimization Diagram forMembrane Separations,” Journal of Membrane Science,129, pp. 19–29 (1997).

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