Advances in Separation & Purification -...

Advances in

Separation & Purification

April 2011

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Contents

Advances in Separation and Purification: mAbs and Beyond

INTRODUCTION

Innovation and the Art of Downstream Processing Uwe GottschalkCreativity may hold the key to advances

in manufacturing technology. s4

DOWNSTREAM CHALLENGES

Addressing the Challenges in Downstream Processing Today and TomorrowGlen R. Bolton , Bernard N. Violand, Richard S. Wright, Shujun Sun, Khurram M. Sunasara, Kathleen Watson, Jonathan L. Coffman, Christopher Gallo, and Ranga Godavarti Newer classes of biotherapeutics will require

innovations in processing technology. s8

ANION EXCHANGE

Benefits of a Revised Approach to Anion Exchange Flow-Through Polish Chromatography Shelly Cote Parra, Christine Gebski A high-performance anion exchange resin performs

well compared with membranes. In addition, the resin

offers greater flexibility and cost savings. s16

VACCINES

Meningitis Vaccine Manufacturing: Fermentation Harvest Procedures Affect Purification Amy Robinson, Shwu-Maan Lee, Bob Kruse, Peifeng Hu Careful analysis of an unusual precipitate is used to

identify its source and correct the manufacturing defect. s21

Cover: Lonza Press Image (2009).

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www.biopharminternational.com April 2011 Supplement to BioPharm International s3

s4 Supplement to BioPharm International April 2011 www.biopharminternational.com

Separation and Purification Introduction

Uwe Gottschalk Phd is vice-president of Purification Technologies

at Sartorius Stedim Biotech GmbH. He is also a member of BioPharm

International’s Editorial Advisory Board and can be reached at uwe.

[email protected].

Innovation and the Art of

Downstream Processing

Creativity may hold the key to advances in manufacturing technology

Uwe Gottschalk

Few people would claim that pro-

cess developers and artists have

much in common. But at the

boundary between art and tech-

nology it is clear that there exists

a shared appreciation of beauty, simplicity

and clarity, in one case generating aes-

thetic appeal, and in the other taking a

complex challenge and crafting from it a

simple and inexpensive solution by imple-

menting more efficient design principles.

Such is the value of innovation in

downstream processing, a concept that

unites the articles presented in this spe-

cial edition of BioPharm International.

Innovation is a difficult concept to define,

and an even more difficult concept to

capture and apply in bioprocessing, par-

ticularly given the industry’s reliance

on long-established technolo-

gies that have evolved incremen-

tally. Process chromatography is

a prime example of incremental

improvement in action. This bed-

rock of biopharmaceutical man-

ufacturing processes has evolved

from a solution to achieve product

quality and therefore regulatory

approvals to a more recent incar-

nation that also focuses on pro-

ductivity and process efficiency.

Unfortunately, this is not a limit-

less resource. Eventually, the steep

hill of improvement flattens out,

and developers are forced to realize that

no more efficiency can be squeezed from

their current processes. Process chroma-

tography has physical limits in terms of

dynamic binding capacity and column

size, above which there are no further cost

savings and therefore no further benefits.

Luckily there is a second form of inno-

vation which is rightly regarded as the

lifeblood of the industry. This involves

the advent of disruptive technologies:

those offering game-changing improve-

ments over a short timescale, initially

serving niche markets and then expand-

ing to challenge the hegemonic position

of established platforms.

Recent examples include disposable

media/buffer bags, the introduction of

membrane chromatography for flow-

through operations, and the replace-

ment of Protein A capture steps with less

expensive alternatives. Only by embrac-

ing such innovations when they occur

can the industry hope to meet the mul-

tiple conflicting demands of a changing

market, with tighter regulations, custom-

ers expecting lower costs of goods, and

increased competition from manufactur-

ers outside Europe and North America.

We are witnessing a revolutionary

change in manufacturing no less impressive

in scale than the revolution in art and lit-


Separation and Purification Introduction

erature during the Renaissance.

This was regarded as a rebirth

of ancient traditions while

acknowledging more modern

developments and applying con-

temporary methods.

Likewise, process develop-

ers are going back to the draw-

ing board to look at ways in

which older technologies can be

merged with new technologies

in more productive ways, in the

context of new development

tools, such as quality by design

and process analytical technol-

ogy. Examples here include the

inclusion of low-cost precipita-

tion steps and high-technology

membrane chromatography

devices in the manufacture of

monoclonal antibodies, the

anc ient and modern com-

bined to reduce the number of

steps and the amount of buf-

fer required, without sacrific-

ing e i ther product iv i ty or

quality, or increasing the foot-

print. Such innovations not

only allow production to keep

up with demand, but allow

new processes to be installed

in existing facilities rather than

requiring upfront investment in

new infrastructure.

This special issue of BioPharm

International presents three arti-

cles from industry leaders pre-

senting a case for innovation

in biomanufacturing. We open

with a thoughtful article by

Ranga Godavarti (Pfizer) asking

whether we really need further

development in downstream

processing or whether incre-

mental improvements in cur-

rent methods will be enough.

Improvements in chromatog-

raphy are discussed by Shelly

Cote (POROS), focusing on pol-

ishing applications for anion

exchange flow-through chroma-

tography. Finally, Shwu-Maan

Lee (Baxter) presents an inter-

esting case study in how pro-

cedures for cell harvesting can

affect the final product.

The art of downstream pro-

cessing is to let innovation create

opportunities for process developers

to explore. In the words of Scott

Adams: “Creativity is allowing

yourself to make mistakes. Art is

knowing which ones to keep.” BP

We are witnessing a

revolutionary change

in manufacturing

no less impressive

in scale than the

revolution in art and

literature during the

Renaissance.

Co

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esy A

uth

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rto

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Separation and Purification Challenges

Addressing the Challenges in

Downstream Processing

Today and Tomorrow

Newer classes of biotherapeutics will require innovations in processing technology

Glen R. Bolton, BeRnaRd n. Violand, RichaRd S. WRiGht, Shujun Sun, KhuRRam m. SunaSaRa, Kathleen WatSon,

jonathan l. coffman, chRiStopheR Gallo, and RanGa GodaVaRti

BeRnaRd n. Violand is a research fellow, Glen R. Bolton is senior principal scientist, RichaRd S. WRiGht is senior principal scientist, Shujun

Sun is senior principal scientist, KhuRRam m. SunaSaRa is senior principal scientist, Kathleen WatSon is associate director, jonathan l.

coffman is associate research fellow, chRiStopheR Gallo is associate research fellow, and RanGa GodaVaRti* is senior director, all from Pfizer. [email protected]

ABSTRACT

In recent years, most pharmaceutical companies

have focused on the development of monoclonal

antibodies (mAbs). Increasing upstream titers and

shrinking development timelines have posed several

challenges to downstream process development

of mAbs. Some of the major strategies and tools

to address these challenges include the develop-

ment of highly efficient platforms, high-throughput

screening (HTS) tools, and reduction of the number

of unit operations to help with facility fit. In the

future, mAbs may represent a smaller percentage

of the pipeline as portfolios concentrate more on

development of antibody fragments, nanobodies,

biosimilar protein therapeutics, conjugated proteins

and vaccines, Fc-fusions, and nonantibody protein

scaffolds. In addition to high cell density mammalian

expression, expanded utilization of other expression

systems such as microbial and yeast will support

these newer biotherapeutics (BioTx). The diversity of

BioTx and expression systems will pose unique chal-

lenges for downstream development. Novel tools,

approaches, and/or platforms will be required to

enable rapid development. Furthermore, with the

increasing emphasis on Quality by Design (QbD),

there is a need to develop paradigms to apply QbD

not only to mAbs but also to other BioTx.

Early biotechnology products in the US

and Europe were characterized by low

titers and low cell densities using vari-

ous cell hosts (1, 2). Most companies had

few recombinant products licensed or

in development, including insulin, somatotropin,

interferon, tissue plasminogen activator, eryth-

ropoietin, Factor VIII, and Factor IX. There were

no benefits or need for platform development

or expanding manufacturing capabilities. Cell

culture was performed using numerous methods,

including roller bottles and stirred tank bioreac-

tors operated in perfusion, batch re-feed, or fed-

batch mode (1).

Sta r t ing in 1986 w ith the l icensu re of

Orthoclone Okt3 (muromonab-CD3), monoclo-



nal antibodies (mAbs) such as

Rituxan (rituximab), Herceptin

(Trastuzumab), and Remicade

(inf liximab) have dominated

the BioTx market (3). Most com-

pany pipelines saw a dramatic

increase in the number of thera-

peutic mAb candidates, which

in turn resulted in a desire to

shor ten development t ime-

lines. Due to the large doses

required, antibodies are typi-

cally expressed at high titers in

high cell density mammalian

cell culture processes. This past

decade has seen both the titers

and cell densities increase by

an order of magnitude to meet

the cl inical and commercial

demands (4–6). Further, com-

panies have made significant

investments in manufactur-

ing facilities, resulting in cell

culture bioreactor capacities of

12,000–20,000L. The increas-

ing volumes, cel l densit ies,

and protein masses have posed

numerous new challenges for

downstream processing.

ADDRESSING CHALLENGES IN mAb

PROCESSING

Shrinking timelines have neces-

sitated developing tools for faster

process development. Because of

the similarity of mAbs, which

mainly differ from each other in

their complementarity determin-

ing regions (CDR), these biomol-

ecules are amenable to platform

development and manufacturing

processes. The standardization

of cellular expression systems,

bioreactor conditions, purifica-

tion processes, manufacturing

hardware, and disposables has

resulted in faster clinical devel-

opment and lower costs. Most

companies have deve loped

standardized purification pro-

cesses typically consisting of a

cell harvest method followed

by a capture chromatographic

step and one or polishing steps.

The approach used at Pfizer has

been described elsewhere (7–12).

Weak partitioning chromatogra-

phy (WPC) and a high through-

put screening (HTS) tools have

been essential for development

of a robust two-column platform

purification process. The benefits

of a two-column purification pro-

cess include reduced capital, foot-

print, quality systems, validation,

cleaning, development costs,

solutions, and water consump-

tion. The HTS method has also

been applied to the development

of other purification processes for

unique biotherapeutic modalities

where screening of a large num-

ber of resins may be required.

Another consideration in mAb

processing is reducing the cost

of goods (COGs). The cost of the

virus-retaining filter, which can

approach that of the protein A

resin, can be reduced by using

prefilters or optimizing protein

concentration, temperature, buf-

fer composition, and solution

pH (see Figure 1) (13). The costs

of freezing and storing bulk

drug substance can be reduced

by targeting high concentra-

tions. Concentrations above 250

g/L have been achieved using

ultrafiltration at elevated tem-

peratures (14) and over 500 g/L

using a wet-ultrafiltration mem-

brane evaporation method (15)

(see Figure 2).

A signif icant consequence

of the t ransfer of products

between facilities is the gen-

erat ion of in-process pools

that exceed the capacity of the

storage vessels. Linking unit

operat ions through tandem

processes would eliminate this

const ra int and addit ional ly

may reduce process time, pro-

cess costs, and documentation

requirements. Feasibility has

been demonstrated for a tan-

dem downstream process for the

purification of mAbs employ-

Figure 1: mAb total mass-throughput and volumetric flux through the Viresolve

Pro filter as a function of time. An X0HC and a Viresolve Pro prefilter were

used with 93 g/L mAb in 3 mM histidine, pH 5.0 at 35 °C.

73

137

192

0

100

200

300

0

50

100

150

200

0 2 4 6 8 10 12 14 16 18 20 22 24

Flu

x [

LM

H]

Tota

l m

ass

-th

rou

gh

pu

t [k

g/m

2]

Time [hr]

Throughput Flux

AL

L F

IGU

RE

S A

RE

CO

UR

TE

SY

OF

TH

E A

UT

HO

RS



ing an affinity Protein A capture

step followed by a flow-through

anion-exchange (AEX) step and

a virus filtration step (VRF)(16).

FUTURE CHALLENGES IN

DOWNSTREAM PROCESSING

Because of the broad range

of new molecules, hosts, and

delivery methods being inves-

tigated, recent biopharmaceuti-

cal development has resembled

that from the early stages of

the industry. A broad range of

a l t e r n at ive B ioTx mo d a l i -

ties have obtained regulatory

approval and are currently in

clinical trials (17–21) includ-

ing antibody fragments, sin-

gle-domain mAbs, Fc-fusions,

vaccines, antibody-drug conju-

gates, nonantibody protein scaf-

folds, PEGylated proteins and

peptides, and viral vectors.

Some of the newer BioTx use

nonmammalian expression sys-

tems such as yeast and E. coli

if there is no requirement for

post-t ranslat ional modif ica-

tions. These alternate expres-

sion systems may offer cost

advantages. However, other

i ssues may ar ise, including

undesired mannosylation in a

transferr in-exendin-4 fusion

molecule expressed in yeast

(22). This unwanted modifica-

tion required an extra process

step using hydrophobic inter-

action chromatography (HIC)

to separate these modified mol-

ecules from the desired product.

E. coli expression often yields

high levels of protein produc-

tion, often in the form of inclu-

sion bodies, which require extra

steps involving their isolation

and the subsequent extraction

and refolding to obtain the

desired product (23).

For a difficult-to-refold pro-

tein such as neurotrophin-4, an

additional step involving sul-

fonation of the cysteines was

also required to obtain a use-

ful refold yield (24). The use of

high throughput screening with

Design of Experiments has been

used to accelerate the determi-

nation of optimal conditions for

refolding difficult proteins (25).

Each of these BioTx modali-

ties poses a unique challenge in

downstream processing because

of the lack of a platform pro-

cess resembling that used for

mAbs. Examples of the purifica-

tion challenges involved with

two Fc-fusion proteins will be

described in the following sec-

tions.

Because each Fc-fusion pro-

tein has a dif ferent protein

sequence fused to an antibody

Fc region, the protein amino

acid sequence, charge, size, and

hydrophobicity vary more than

those of antibodies. In addi-

tion, they are often thermally,

chemically, or enzymatically

unstable, which can lead to

high levels of aggregates, clips,

or inactive species.

Fc-FUSION PROTEINS

Case study 1.

A pur i f icat ion process was

d e v e l o p e d f o r a n a c i d i c

Fc- f usion prote in that was

unstable at ambient tempera-

tures and low pH and that

quickly formed high molecular

weight oligomers (HMW1) and

dimers (HMW2).

To improve stability, the pro-

tein A chromatography step was

performed at 2–8 oC and the

low pH eluate was neutralized

immediately. The capacity of

the protein A resin for the mol-

ecule was less than 15 mg/mL

which is typical for Fc-fusion

proteins.

A combination of HTS meth-

ods and gradient chromatogra-

phy was employed to identify

resins and operating conditions

for h igh molec u la r we ight

(HMW) species removal. Anion

Figure 2: Viscosity versus concentration for an Fc-fusion protein and a mAb

concentrated using a wet-ultraflltration membrane evaporation method.

0

Fc-Fusion protein

mAb

1

10

Vis

cosi

ty [

cP]

100

1,000

10,000

100,000

100 200

Concentration [g/L]

300 400 500



exchange, cation exchange, and

HIC resins were screened in 96

well plates in an HTS format.

Additionally, ceramic hydroxy-

apat ite (CHA), immobi l i zed

metal affinity chromatography

(IMAC), and HIC methods were

screened with elution gradients.

Due to the low pI of this

Fc-fusion protein, a WPC–AEX

step was not successful. Based

on the resin screening, CHA

chromatography was selected

for optimization. An initial test

of CHA was performed using

a phosphate g rad ient . This

method allowed identification

of step elution conditions that

reduced the HMW levels from

29% to 10.1% with 87% yield.

The HMW1 f lowed through

the column and the peak con-

tained primarily dimer species

(HMW2). The CHA step had a

capacity of 20 g/L.

The HTS analyses also indi-

cated that a HIC resin under

certain loading and elution con-

ditions could provide reduction

of HMW. An initial experiment

using protein A peak pool and a

salt gradient indicated the resin

had a relatively low capacity for

the product but was effective at

reducing both the levels HMW

species and less active monomer

variants.

Fur ther development was

performed using the protein A

pool as a load to the HIC step to

identify conditions where the

product would f low through

the column while the HMW

and less active monomer would

bind. A successful method was

developed but the HMW1 spe-

cies began to flow through the

column at about 10 mg/mL of

load.

Because the CHA step was

successf u l at remov ing the

HMW1, but not HMW2 species,

the CHA pool was used as a load

for the HIC step. The HIC step

was loaded with the CHA pool

containing primarily HMW2 at

20 mg/ml. The HIC step pro-

vided a six-fold reduction in

HMW2 levels. In addition, the

HIC step reduced the level of

host cell proteins by 1.4 log10.

Based on these data, the final

process used a Protein A cap-

ture step with elution at low pH

followed by immediate neutral-

ization. This was followed by a

bind-elute CHA step to remove

H M W1 fol lowed by a f low

through HIC step to remove

HMW2 (see Figure 3). The first

two chromatography steps were

performed in the cold to main-

tain product stability.

Case study 2.

Dur ing deve lopment of an

early-phase purification process

for an acidic Fc-fusion protein

expressed in a CHO cell culture

system, it was determined that

the protein displayed a range

of 24–100% activity. Several

types of chromatographic res-

ins, including cation exchange,

anion exchange, HIC, CHA,

and IMAC, were evaluated to

remove the less active mole-

cules along with other product

and process-related impurities.

The process and condit ions

described above for Fc-fusion

protein 1 were not applicable to

Fc fusion protein 2.

The CHA is a bimodal resin

with two primary functional

groups: phosphate and Ca 2+.

The negatively charged phos-

phate groups serve as cation

exchangers and the Ca2+ groups

serve as chelators to proteins.

For purif ication of basic and

neutral proteins, the CHA col-

umn is usual ly equi l ibrated

with phosphate buf fer at a

neutral pH to utilize a cation

exchange mechanism for the

separation of protein monomers

from HMW species.

T he CH A colu m n, when

equilibrated with phosphate,

Figure 3: Yield and percentages of HMW1 and HMW2 versus step in the

Fc-fusion puriflcation process described in case study 1.

Protein A

HMW1

HMW2

Yield

0

10

20

30

40

50

60

70

80

90

100

Perc

en

t

Bind-elute CHA Flow-through HIC



r e solve d f u l ly ac t ive pro -

tein from partially active pro-

tein. The product was in the

unbound fraction while the less

active molecules were bound to

the column. However, the load

capacity was low and removal

of high molecular weight spe-

cies was poor using the phos-

phate-charged CHA column in

flow-through mode.

A calcium-charged CHA col-

umn operated in bind-elute

mode successf u l ly removed

the less active monomer spe-

cies, as well as high molecular

weight species, from the process

stream. In this case, the CHA

column was equilibrated with a

low concentration of CaCl2 in

a neutral pH buffer, and then

loaded with a protein mixture

also containing CaCl2. After

washing the column with a

neutral pH buffer, the column

was eluted with a phosphate

buffer and stripped with a high

phosphate buffer. A high yield

of the product was obtained

in the elution peak. The less

active protein and HMW species

were present in the strip frac-

tion. These data demonstrate

that although Fc-fusion proteins

are more challenging to purify

than mAbs, other chromatogra-

phy resins can be used to yield

high-quality product.

These two examples illustrate

the challenges posed by the dif-

ferent Fc-fusion molecules. In

one case, calcium charging of

the CHA column was required

for removal of a lower activity

monomer species. In another

case, a bind-elute CHA column

fol lowed by a f low through

HIC column was required for

adequate HMW removal, resin

capacity, and product y ield.

The sequence of the polishing

steps is also important for both

purity and capacity.

ANTIBODY-DRUG CONJUGATES

Numerous antibody drug conju-

gates (ADCs) are in clinical trials

with most of them being used

in cancer treatment with a toxin

payload (26). The major down-

stream challenge is separation of

multiple species containing dif-

ferent levels of the attached drug

to the mAb. This results from

the limited specificity for chem-

ical conjugation of the drug to

either exposed lysines or cyste-

ines (generated by limited reduc-

tion) on the mAb. To enable

straightforward downstream pro-

cessing of ADCs, two main strate-

gies have been used to generate

more homogeneous molecules.

These are the incorporation of

additional free cysteines that can

then be chemically targeted spe-

cifically or to mutate out some

of the native cysteines to serines

so that they are less available for

coupling after a limited reduction

reaction (27, 28).

mAB FRAGMENTS AND SINGLE

DOMAIN mAbs

One approach to generate smaller

and simpler BioTx for down-

stream processing is to replace

full-length mAbs with anti-

body fragments such as single-

chain variable fragments (scFvs)

and antigen-binding fragments

(Fabs). Cimzia and Lucentis are

two recently approved Fabs pro-

duced in E. coli (29) and although

no scFv has yet received regula-

tory approved, there are a large

number of both types of Ab frag-

ments in clinical trials (30–32).

Purification of these Ab frag-

ments utilizes normal chromato-

graphic procedures unless the

H-chain is from the VH3 IgG

gene family in which case it will

bind to protein A. After purifica-

tion these small fragments are

usually modified to improve their

pharmacokinetics using tech-

nologies, such as PEGylation.

PEGylation dramatically changes

the chromatographic properties

of proteins such that loading

capacity is severely compromised

thus requiring larger columns

(33, 34). Additionally PEGylation

can dramatically increase the vis-

cosity of concentrated drug sub-

stances.

In addit ion to mAb f rag-

ments, another approach to

obtain smaller less complex

antibody molecules is to use

either Domain Abs (human

derived) or nanobodies (Llama

derived), which are composed

of a single var iable domain

of ~14kDa. These small single

domain antibodies can be eas-

ily captured from fermentation

broth using Protein A since

they can be derived from the

VH3 IgG gene family (35, 36).

Subsequent downstream purifi-

cation has been accomplished

u s i ng s t a nd a rd c h romato -

graphic techniques (36).

VIRAL VECTORS

Using viral vectors as BioTx for

gene therapy offers many excit-

One approach to

generate smaller

and simpler BioTx for

downstream

processing is to

replace full-length

mAbs with antibody

fragments



ing opportunities to inactivate

new ta rget s . T he i r produc-

tion presents numerous chal-

lenges because of the complex

nature of their composit ion

and large size (2–5 x 106 kDa).

They are usually composed of

multiple proteins that encap-

sulate the nucleic ac id pay-

load. Additionally, there may

be empty capsids that do not

contain the gene of interest but

are similar in physicochemical

properties and can be difficult

to separate f rom the desired

product. CsCl density gradient

centrifugation has long been

the method of choice for their

isolation; however, for clinical

test ing, larger scale produc-

tion has recently been accom-

plished using standard column

ch romatog raph ic met hods .

However, because of their large

size, their binding capacity

is usually very low and chro-

matographic media with nor-

mal pore sizes used for proteins

offer no advantages since the

pores are not available. Smaller

beads with more surface area

is the most effective manner

to increase capacity (37,38).

Other processes that are being

utilized include tangential flow

filtration that takes advantage

of the large part icle size for

separation from smaller cellu-

lar contaminants.

VACCINES

Vacc i nes ca n be produced

u s i ng nu merous te ch nolo -

gies and have been generated

against proteins, polysaccha-

ride, and small molecule tar-

gets. Pneumonia is the largest

single infectious disease caus-

ing infant mortality (39). To

date, three vaccines have been

licensed for treating diseases

caused by pneumococcus-con-

taining polysaccharides conju-

gated to proteins (40).

Because prophylactic vaccines

are given to a large healthy pop-

ulation ranging from infants

to the elderly, the regulatory

requirements during develop-

ment and for licensure tend to

be more stringent as compared

with other biotherapeutic prod-

ucts. This also results in more

early investment and upfront

development work for a vaccine

candidate than for a model bio-

therapeutic agent. Most com-

mercial prophylactic vaccines

a re not we l l - cha rac te r i zed

biologicals from a regulatory

perspect ive because of their

inherent complex ity and/or

poorly understood mechanism

of action. In addition, if an in

vivo animal potency model is

available, the results don’t usu-

ally predict the human response

to the product with regards to

the protein structure-immu-

nogenecity/ant igencity rela-

tionship. These reasons force

sponsors to initiate extensive

process characterization stud-

ies and to lock the production

processes early in the clinical

development program (41).

One new approach to gen-

erate vaccines is to use virus–

like particles (VLP) as potent

immunostimulaters with cova-

lent attachment of many copies

of the desired antigen to their

surface (42). Similar to viral

vectors, VLPs present unique

chal lenges because of thei r

large size (2.5 x106 kDa). Qß

(14 kDa) is the most commonly

used VLP. It is derived from the

structural coat protein of this

virus and it naturally assembles

into 180 copies of itself during

expression in the cytoplasm

of E. coli. Subsequent purifica-

tion is accomplished by clas-

sical column chromatography

techniques with size-exclusion

chromatography cited most fre-

quently, which is not a desired

step for scale-up (41). The chal-

lenges to develop alternative

process steps are similar to

the viral vector BioTx because

of low capacity for most resins

as well as the fact that there

may be heterogeneous levels of

antigen attached to the VLPs.

Another new approach is the

del ivery of plasmid DNA to

cells, which leads to the tran-

scription of antigens and a sub-

sequent immune response (43).

An addit ional complex ity

is the fact that many vaccines

are multivalent. The manufac-

turing processes for these vac-

cines involve many different

fermentation, purification, and

conjugation trains for the poly-

saccharides and carrier proteins.

Developing processes that are

similar between antigens can

improve manufacturability and

facility fit (44).

BIOSIMILARS

T h e F D A a n d E u r o p e a n

M e d i c i n e s A g e n c y ( E M A)

have indicated that the clini-

cal testing requirements for a

biosimilar drug can be reduced

if it can be demonstrated that

the biosimilar candidate does

not meaningfully differ from

the innovator drug (45, 46).

Currently developed biosimi-

lars, such as human growth

hormone, insulin, and erythro-

poietin, are chemically far sim-

pler than mAbs. It is unlikely

that blood or plasma-derived

products or complex vaccines

will be acceptable as biosimi-

lars in the US or Europe due to

their complexity (47). However,

eight categories of biosimilar

molecules have been manu-



factured in India, including

monoclona l ant ibodies and

recombinant hepatitis B vac-

cine (48). The first monoclonal

antibodies that will lose patent

protection in Europe and the

US were made using more strin-

gent impurity level targets than

exist now because the technol-

ogy is more mature and the

safety and efficacy is well estab-

lished. For example, early mAbs

had levels of high molecular

weight species below 0.5% (49).

It is possible that mAbs in the

future will target levels below

5% (50).

A new purification challenge

will be to produce antibodies

with product quality attributes

that match the innovator anti-

bodies. This may be carried out

with a host cell and/or purifica-

tion process that differs from

the innovators. A biosimilar

antibody can have dif ferent

levels of impurity or product

variant species than the inno-

vator, i f it can be just i f ied.

The biosimilar manufacturer

will determine the amount of

clinical and characterization

data that will be used to jus-

t ify differences to the agen-

cies. Therefore each biosimilar

manufacturer must balance the

development costs required to

match innovator antibody char-

acteristics against the costs of

justifying differences. During

deve lopment of b ios im i la r

recombinant human growth

hormone, higher levels of host

cell proteins lead to increased

levels of antibodies to both

host cell proteins and growth

hormone. This necessitated

modifications to the purifica-

tion process, which eventually

allowed demonstration of com-

parability to the innovator and

approval by the EMA (51). The

success of a biosimilar devel-

opment program will benefit

from robust, selective, and effi-

cient purification processes that

require minimal development

time and material.

IMPLEMENTING QUALITY

BY DESIGN

The past decade has witnessed

an increasing emphasis—by

both industry and regulatory

agenc ies— on implement ing

Quality by Design (QbD) prin-

ciples. In 2008, the FDA’s Office

of B iotechnolog y P roduc t s

invited companies to partici-

pate in a pilot program involv-

ing the submission of quality

information for biotechnology

p ro duc t s i n a n E x pa nde d

Change Protocol (52,53). The

purpose of the pilot program is

“to gain more information on

and facilitate agency review of

QbD, risk-based approaches for

manufacturing biotechnology

products”. These approaches

link “attributes and processes

to product performance, safety

and efficacy”. The concept of

a design space has been intro-

duced, which is defined as the

multidimensional combination

and interaction of input vari-

ables and process parameters

that have been shown to pro-

vide assurance of quality (54).

The underlying principles of

QbD and risk management are

contained in ICH Q8 (R2), Q9,

and Q10 (50, 54–56).

A comprehensive case study

of the application of QbD prin-

ciples to the development of

a mAb product (A-mAb) has

recently been published that

describes extensive use of scale-

down models to gain process

and product knowledge (50).

Howeve r, s i g n i f i c a nt d at a

from scale-down models could

be used to define the design

space. In addition, data from

triplicate studies performed at

pilot or commercial scale using

center-point conditions could

ensure that product quality and

process performance were not

impacted by the process change.

Signi f icant quest ions and

challenges remain in imple-

menting QbD. While there is

agreement on the value of pro-

cess and product understand-

ing, there is often debate within

most biopharmaceutical com-

panies on the extent of QbD

investment and underlying cost

implications. Several technical

hurdles need to be overcome

to def ine methodolog ies to

describe a design space and con-

trol strategy. Quality systems

need to be developed that will

enable the movement within a

design space. Furthermore, as

different regulatory agencies

embrace QbD along different

timelines, how will that affect a

global submission? Finally how

will QbD apply to more chal-

lenging large molecules such

as vaccines? People involved

in process development will be

involved in the effort required

to resolve many of these ques-

tions.

CONCLUSION

Biopharmaceut ica l develop -

m e nt a n d m a nu f a c t u r i n g

have evolved significantly in

the past 25 years. Clinical and

commerc ia l pipe l ines have

evolved from replacement pro-

teins and other therapeut ic

proteins/hormones to mAbs.

Tech nolog ica l adva nces i n

bioprocessing have led to tre-

mendous increases in prod-

uc t a nd ce l l mass t hat in

turn have posed several chal-

lenges to downstream process



development of mAbs. Novel

approaches to downstream pro-

cessing, such as development

of highly efficient platforms,

HTS tools, and reduction in the

number of unit operations have

helped to address these chal-

lenges. In the future, mAbs may

represent a smaller percentage

of the pipelines and the chal-

lenges to downstream process-

ing will likely have a different

focus. Most product portfolios

will likely include other com-

plex biomolecu les, such as

conjugated proteins and vac-

cines, a variety of biosimilar

protein therapeutics, and novel

scaffolds such as fusion pro-

teins, nanobodies, etc. Future

expression systems may include

microbial, yeast, and others in

addition to high cell density

mammalian systems. The diver-

sity of scaffolds and expression

systems will pose unique chal-

lenges for downstream develop-

ment. Novel tools, approaches

and/or plat forms may need

to be applied to enable rapid

deve lopment . Fu r ther more,

there is a need to develop para-

digms to apply QbD not only

to mAbs but also to other large

molecules. BP

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Separation and Purification Anion Exchange

Benefits of a Revised Approach

to Anion Exchange Flow-Through

Polish Chromatography

A high-performance anion exchange resin performs well compared with membranes. In addition, the resin offers greater

flexibility and cost savings.

Shelly Cote Parra, ChriStine GebSki

Shelly Cote Parra, MS, is a senior field applications scientist in the POROS Applications group at Life Technologies, [email protected]. ChriStine GebSki, MS, is the director of POROS R&D and Applications at Life Technologies, Bedford, MA, [email protected].

AbStrAct

Anion exchange (AEX) products are commonly used as

a polish step in product flow-through (Ft) mode to bind

impurities. comparing resins with membranes and mem-

branes to membranes is challenging due to the com-

plicated and unique scale-down model formats of the

products offered by different vendors. A novel approach to

AEX Ft using a short, 5 cm length, packed bed format with

a faster operating flow rate is explained. the performance

of commonly used AEX resins and membrane adsorbers,

detailing dynamic binding capacity performance, effi-

ciency, and a new disposable option for packed bed chro-

matography are compared. the data presented will show

that a high performance AEX resin competes well with the

performance of membranes. It provides similar processing

times and the added benefits of reusability, ease of packing

at different scales in various column formats, and the ability

to implement initial process design from early phase manu-

facturing to commercial manufacturing, reducing overall

costs and time to market.

AEX chromatography resins and mem-

brane adsorbers are frequently used for

downstream purification in the biotech-

nology industry. AEX products are used

for polish chromatography in product FT

mode to bind product-related and process impurities.

Multiple product formats are commercially available,

including chromatography resins, which are packed

into chromatography columns, and membrane adsorb-

ers, which are supplied in self-contained plastic hous-

ings. Comparing performance across product types

and formats, including resins to membranes and mem-

branes to membranes, is challenging due to the compli-

cated and unique formats for scale-down models of the

different product types.

In FT mode, membranes have shown advantages over

traditional soft gel packed beds due to faster operating

flow rates, reduced buffer requirements, and dispos-

ability. In most cases, traditional soft-gel FT columns

are sized for the optimization of volumetric throughput

to improve operating flow rate and decrease process

bottlenecks, rather than sized for actual impurity bind-

ing capacity. This results in packed columns with larger

diameter, and therefore volume, than optimal, and this

improper column sizing results in greater resin require-

ments and increased buffer usage, both of which impact

operating time and cost of goods. Although membranes

may be simpler to implement for early phase manu-

facturing, where the process scale is typically smaller,

these products are not always cost effective at the larger,

late phase or commercial manufacturing scales due

to high material costs and lack of reusability. In addi-



* The operating linear flow rate was decreased due to increased pressure and cartridge leaking during the study

AEX ProductColumn/

Membrane Volume (ml)

Column/Membrane DimensionsLinear Flow

Rate(cm/h)

Residence Time (min)

POROS HQ 50, Life Technologies 0.83 0.46 cmD x 5 cmL 1000 0.3

Fractogel TMAE (M),EMD Millipore

0.83 0.46 cmD x 5 cmL 800 0.4

Q Sepharose Fast Flow, GE Healthcare

0.83 0.46 cmD x 5 cmL 500 0.6

Sartobind Q Nano, Sartorius 11.75 cm x 0.4 cm

15 membrane layers250 DNA100 BSA*

0.1 DNA, 0.3 BSA*(Target 10 MV/min)

Mustang Q Coin, Pall 0.351.43 cm x 0.2 cm

16 membrane layers131

0.1 (Target 10 MV/min)

Table I: Scale down approach for AEX chromatography resin and membrane adsorber comparison.

tion, the available product formats

are limited, posing linear scale-up

challenges, adding to the challenge

of maintaining process continuity

and proving process equivalence

between scales. Late-phase process

redesign or a complete switch to

packed bed chromatography may

be required due to high consumable

costs or limited product formats not

allowing for linear scale-up.

On the other hand, the re-use

of resins is well established for

traditional packed bed chroma-

tography and this facilitates lower

overall material costs. In addi-

tion, chromatography columns

are easily scalable. Unit opera-

tions can be defined and locked

for early phase manufacturing

and then simply scaled up in a

linear fashion as manufacturing

scale increases reducing the need

for later phase redesign. High

performance, rigid resins with

a higher volumetric throughput

capability provide an advan-

tage over soft-gel resins allow-

ing for properly sized columns

with smaller footprints similar

to membranes. These rigid resins

allow for convective flow through

the bead, improving mass trans-

fer and increasing efficiency at

higher linear flow rates, thereby

improving process productivity.

Demonstration of performance

using a scale-down model

An alternative and beneficial

approach to AEX FT is the use of

a short, 5-cm length, packed bed

format with a faster operating flow

rate. This format enables volumetric

throughput capability that is simi-

lar to the membrane format, and

increases flexibility when designing

a purification scheme. The scale-

down model used to compare the

performance of the different AEX

products/formats is summarized

in Table I. We evaluated five com-

monly used AEX products: three

resins and two membrane adsorb-

ers for dynamic binding capacity

of DNA and protein using bovine

serum albumin (BSA) to mimic

common contaminants, such as

host cell proteins, removed during

a FT step. The resin target operating

flow rate was the maximum flow

rate defined per the manufacturer’s

operating instructions. The mem-

brane target operating flow rate was

based on the common industry

design space of 10 membrane vol-

umes (MV) per minute. Although

POROS chromatography resins can

be operated at 2000 cm/h or faster,

1000 cm/h was evaluated as the

upper limit for this evaluation. A 5

cmL POROS column can be oper-

ated at 1000 cm/h with a low pres-

sure drop allowing for the use of

high operating flow rates in con-

ventional low pressure chromatog-

raphy columns and systems (see

Figure 1).

Materials and methods

DNA Dynamic Binding Capacity:

The AEX product was pre-charged

with 20 mM sodium phosphate,

1 M NaCl, pH 7.0 followed by

an equilibration with 20 mM

sodium phosphate, 50 mM NaCl,

pH 7.0 (7.8 mS/cm). Each col-

umn/membrane was loaded with

2 mg/mL Herring Sperm DNA

(Sigma D3159) in equilibration buf-

fer (titrated to pH 7.0 with 0.2 M

sodium phosphate dibasic, anhy-

Products are not

always cost-effective

at the larger, late

phase or commercial

manufacturing scales

due to high

material costs and

lack of reusability.

AL

L F

IGU

RE

S A

RE

CO

UR

TE

SY

OF

TH

E A

UT

HO

RS



Figure 1: Pressure-flow curve for POROS HQ 50 in 8 cmD Go-Pure

Pre-Packed column format at bed heights of 5 cm and 20 cm (5 um frits, 0.1 M

NaCl, system pressure subtracted).

Flow rate (cm/hr)

0

0.0

1.0

2.0

3.0

4.0

5.0

5 cm Bed height 20 cm Bed height

200 400 600 800 1000

Pre

ssu

re (

ba

r)

A shorter column

run at a faster

operational flow rate

can achieve good

viral and impurity

clearance.

Figure 2: Effect of load conductivity on viral clearance capability of POROS HQ.

Conductivity (mS/cm)

0

0

1

2

3

4

5

6

XMuLV MVM

5 10 15 20

Vir

al

cle

ara

nce

(Lo

g1

0)

drous) with a final conductivity

of 8.8 mS/cm per the flow rates

listed in Table I. Binding capacities

at 5% (C5) and 50% (C50) break-

through were determined based

on UV absorbance.

BSA Dynamic Binding Capacity:

The AEX product was pre-charged

with 20 mM Tris, 1 M NaCl, pH

8.0 followed by an equilibration

with 20 mM Tris, pH 8.0 (1.1 mS/

cm). Each column/membrane was

loaded with 10 mg/mL BSA (Sigma

A7906, pI 4.7-5.3, MW 66 kDa) in

equilibration buffer with a final

conductivity of <2 mS/cm per the

flow rates listed in Table I. C5 and

C50 breakthrough were determined

based on UV absorbance.

POROS HQ Viral Clearance:

Polyclonal human IgG (Sigma

G4386, MW:155–160 kDa; pI: ~6.9)

was used for the model process. The

salt concentrations evaluated and

the corresponding conductivity val-

ues are summarized in Table III and

Figure 2. The column format was

0.46 cmD x 5 cmL, 0.83 mL or 0.46

cmD x 20 cmL, 3.3 mL. Viral clear-

ance was assessed at 1000 cm/hr at

room temperature for the 25 mM,

50 mM and 150 mM runs and 300

cm/hr for the 100 mM NaCl run.

The studies were all run at pH 7.0

using 20 mM bis-tris propane for

buffering. The column was loaded

with 5 mg/mL IgG with a 5%

xenotropic murine leukemia virus

(xMuLV, retrovirus, enveloped,

ssRNA, 80-120 nm) or murine min-

ute virus (MVM, parvovirus, non-

enveloped, ssDNA, 18-26 nm). Spike

and column FT samples were taken

to determine viral clearance.

Results and discussion

The DNA and BSA binding capac-

ity data of five anion exchangers

is summarized in Table II. POROS

HQ 50 demonstrated the highest

DNA binding capacity of the five

AEX products assessed. The C5/

C50 ratio is a measure of the mass

transfer capability of the resin/

membrane and a way to character-



ize the efficiency of binding. The

more efficient the mass transfer

capability of a product, the less

effect flow rate has on capacity. As

the C5/C50 ratio approaches 1.0

the resin/membrane becomes more

efficient. The POROS HQ C5/C50

ratio is as efficient as the mem-

branes and significantly more effi-

cient than Fractogel TMAE and Q

Sepharose FF for DNA binding. The

capacity at 5% breakthrough on

the 5 cmL column at 1000 cm/hr

was compared to the capacity on a

20 cmL column run at 300 cm/hr.

POROS HQ had minimal change

in performance in the two formats

as compared to the other two res-

ins. In addition, with POROS HQ,

the DNA capacity is high under

a wide range of operating condi-

tions (30-35 mg/mL at pH 6.0–9.5

with 150 mM NaCl and >20 mg/

mL at pH 7.0 with up to 400 mM

NaCl, data not shown). POROS

HQ ranked second highest for BSA

C5 dynamic binding capacity and

shows similar capacity and mass

transfer efficiency similar to the

membrane products.

POROS HQ in FT mode demon-

strated good viral clearance capabil-

ity for XMuLV up to 150 mM NaCl

(18 mS/cm) at pH 7.0, as summa-

rized in Table III and Figure 2. The

MVM model virus showed good

clearance up to 50 mM NaCl (8 mS/

cm) in this new AEX FT format

suggesting that a shorter column

run at a faster operational flow rate

can achieve good viral and impu-

rity clearance. The conductivity of

the load appears to have an effect

on both viruses. MVM is a poorly

charged virus so minimal salt is

needed to neutralize the charge and

decrease the binding. However, it is

a small virus and can easily access

the pores, so binding performance

is flow rate independent. XMuLV,

on the other hand, is significantly

larger and highly charged. With the

higher salt (18 mS/cm), the hydro-

dynamic radius of the virus is most

likely changing, allowing for more

optimal perfusion into the bead.

Table IV presents a cost model

comparing POROS HQ 50 in this

short-bed format to a traditional

resin and membrane process. The

POROS HQ product load time in

the new format is seven times faster

than the traditional resin step and

three times faster than the mem-

brane. The total process time is six

times faster than the traditional

resin process and almost two times

faster than the membrane. In addi-

tion, the optimized HQ format uses

four times less buffer than the tradi-

tional resin step. One of the benefits

of using a resin is the reusability

at commerical scale, and POROS

HQ allows for aggressive cleaning

and sanitization, yielding excellent

cycling and reuse performance.

This study shows the cost difference

for one cycle compared to 50 cycles

and the cost benefit of reuse com-

pared to a single-use membrane.

Conclusion

The novel approach of using a

short, disk-like column with

a resin capable of operating at

high flow rates for AEX flow-

through polish chromatogra-

phy delivers increased flexibility

when designing a purification

scheme. This format is ideal for

rigid resins with flow rate inde-

pendent performance driven

by the convective properties of

the base bead. The properties of

POROS HQ, for example, drive

increased throughput and smaller

column sizes, and ultimately a

column that can be sized based

on capacity for impurities. In

addition, POROS HQ has been

DNA Binding Capacity BSA Binding Capacity

AEX Product

C5(mg/ml)

C50(mg/ml)

C5/C50Ratio

Δ in C5 from C5 at 300 cm/hr 20 cmL

C5(mg/ml)

C50(mg/ml)

C5/C50Ratio

Δ in C5 from C5 at 300 cm/hr 20 cmL

POROS HQ 50 37 52 0.72 5% 77 103 0.69 16%

Mustang Q 33 48 0.70 n/a 86 99 0.84 n/a

Sartobind Q Nano 18 25 0.72 n/a 49 69 0.65 n/a

Fractogel TMAE 10 26 0.38 53% 46 60 0.72 49%

Q Sepharose Fast Flow 6 43 0.13 64% 20 41 0.45 59%

Table II: DNA and BSA dynamic binding capacities of flve AEX products.

Virus Load NaCl Concentration

(mM)

Load Conductivity (mS/cm)

Viral Clearance (Log 10)

XMuLV MVM

25 5 4.04 4.98

50 8 4.95 4.34

100* 13 >5.13 2.30

150 18 >5.26 0.83

Table III: Viral clearance on POROS HQ (1000 cm/hr, 5-cm bed height).

*100 mM NaCl run was executed using a 20-cm bed height column at 300 cm/h.



* Total process time includes a 30 minute sanitization hold time for chromatography resin reuse.

Table IV: Large scale operating cost model for POROS HQ versus conventional AEX resin (2700 L at 5 mg/mL, 13.5 kg

monoclonal antibody process).

POROS HQ in Optimized Format

Conventional Soft-Gel Resin

AEX Membrane

Column/Membrane Dimensions 80 x 5 cm 80 x 20 cm Jumbo

Column/Membrane Volume (L) 25.1 100.5 5.0

Load Capacity (mg of protein/mL of resin) 538 134 2,700

Linear Flow Rate (cm/hr or MV/min) 1000 150 5

Product Load Process Time (h) 0.5 3.6 1.8

Total Process Time (h)* 1.1 6.7 1.9

Buffer Volume (L) 602 2412 252

Pre-Equil: 3 CV 3 CV 10 MV

Equil: 5 CV 5 CV 10 MV

Wash: 3 CV 3 CV 10 MV

Regeneration: 3 CV 3 CV N/A

Sanitization: 3 CV 3 CV N/A

Storage: 3 CV 3 CV N/A

Buffer Cost ($) 1807 7236 756

Column Packing Labor Cost ($) 4500 4500 0

Process Labor Costs ($) 3937 5499 2382

Cost of Resin/Membrane ($) 50,200 100,500 78,000

Total Cost of Processing/Cycle ($)

1 Cycle 60,444 117,735 81,138

5 Cycles 16,684 33,735 N/A

10 Cycles 11,214 23,235 N/A

20 Cycles 8,479 17,985 N/A

50 Cycles 6,838 14,835 N/A

shown to have high impurity

binding capacity and clearance

over a range of process condi-

tions, including high conductiv-

ity conditions (data not shown).

This decreases the need for dilu-

tion of the feed stream or inclu-

sion of a diafiltration step prior

to loading on the HQ column,

making the process more efficient

and cost-effective. If disposabil-

ity is a factor, Life Technologies

now offers Go-Pure Pre-Packed

Chromatography Columns for

maximum convenience, provid-

ing faster time to process and

faster time between processes.

This new approach to AEX FT

polish chromatography increases

process development flexibility

and offers a more cost effective

approach to this process step as

compared to improperly, oversized

soft gel columns, and membranes

with limited sizes and expensive

formats at larger scale. This study

shows that a high performance

AEX resin competes well with the

performance of membranes, provid-

ing similar processing times and

the added benefits of reusability,

which decreases material costs with

increased cycles as modeled in Table

4. In addition, ease of packing at

different scales in various column

formats and the ability to imple-

ment initial process design from

early phase manufacturing to com-

mercial manufacturing reduces

overall process development costs

and decreases time to market.

For Research Use Only. Not

intended for animal or human

therapeutic or diagnostic use.

Acknowledgments

The authors would like to thank

Su s a n ne A le x a nde r, Roge r

Decker, and Elliot Haimes for

assistance with the execution of

these studies. BP


Separation and Purification Vaccines

Meningitis Vaccine Manufacturing:

Fermentation Harvest Procedures

Affect Purification

Careful analysis of an unusual precipitate is used to identify its source and correct the manufacturing defect.

Amy Robinson, shwu-mAAn Lee, bob KRuse, Peifeng hu

Amy Robinson, PhD is a senior manager, shwu-mAAn Lee,* PhD is a technical director, bob KRuse, PhD is a research scientist, and Peifeng hu,

PhD is a principal scientist, all at Baxter Healthcare, 8000 Virginia Manor Road, Suite 140. Beltsville, MD 20705. [email protected].

Ph

oto

co

urt

esy o

f th

e a

uth

ors

AbStrAct

the meningitis vaccine NeisVac-c is a group c menin-

gococcal polysaccharide conjugated to tetanus toxoid.

the polysaccharide is recovered from the Neisseria

meningitidis cell capsule and is purified by base treat-

ment with subsequent diafiltration to remove hydro-

lyzed cell impurities. Purified polysaccharide is clear to

slightly cloudy. A recent group of successive lots con-

tained large amounts of precipitate that had not been

observed in 10 previous years of commercial manu-

facturing. the precipitate was mostly composed of the

sodium salt of palmitic acid (c16:0) with lesser amounts

of palmitoleic (c16:1), oleic (c18:1), stearic (c18:0) and

myristic (c14:0) sodium salts. the elevated fatty acid

levels that formed the precipitate were linked to a dam-

aged pump used during harvest. replacement of the

damaged pump corrected the issue and >15 lots have

since been produced without precipitation.



Figure 1: Visible precipitation during GCMP purification.

N eisVac-C is a vac-

cine that prevents

the invasive dis-

e a s e c au s e d by

Neisse r ia menin -

gitidis serogroup C. The active

ingredient is a polysaccharide–

protein conjugate. Each dose

contains 10 µg of de-O -acety-

lated group C meningococcal

polysacchar ide (GCMP) con-

jugated to 10-20 µg of tetanus

toxoid protein and adsorbed

onto 0.5 mg of aluminum as

aluminum hydroxide in saline.

The GCMP is isolated from

the culture medium of Neisseria

meningitidis by microfiltration.

The microfiltration permeate is

concentrated and diafiltered to

remove small soluble fermen-

tation components. The major

purification step is a saponifi-

cation reaction in which the

GCMP is ref luxed with high

concentrations of base for sev-

eral hours. After diafiltration,

the mixture typically appears

clear to slightly cloudy. A recent

cohort of successive GCMP lots

contained a precipitate that had

not previously occurred in ten

years of commercial manufac-

turing (see Figure 1). This article

describes the problem, the root

cause analysis, and the correc-

tive actions.

Materials and methods

The antigenic component of

the vacc ine, the de - O -acet-

ylated GCMP, was pur i f ied

from the culture supernatant

of Neisseria meningitidis sero-

group C, strain C11. The cells

and spent culture medium were

circulated through 0.2 µm hol-

low fiber cartridges using a cir-

cumferential piston pump. The

permeate from the f i ltration

contained the polysaccharide

(see Figure 2).

This filtrate was concentrated

and diafiltered across a 300 kilo-

dalton (kDa) nominal molecu-

lar weight cutoff (NMWCO)

u lt ra f i lte r ( U F ), ( Mi l l ipore

Pellicon 2) (see Figure 2, step

4), which retains the GCMP.

The concentrated GCMP was

then chemically modified with

a saponif icat ion react ion to

remove acetyl groups (Figure 2,

step 5). Base treatment is the

major purif ication step with

high temperature incubation for

several hours in NaOH.

Deacetylat ion removed al l

the acetyl g roups f rom the

O-positions and most of the ace-

tyl groups from the N-positions

of the GCMP. It also hydrolyzed

cell impurities and saponified

any fatty acids, which were

removed by subsequent diafil-

tration with water for injection

(WFI) across a 50 kDa NMWCO

UF (Pellicon 2, Millipore) (see

Figure 2, step 6). The GCMP

remain in the UF retentate.

T he N - ace t y l g roups a re

believed to be immunologically

important, and were restored

in a subsequent chemical reac-

tion (see Figure 2, step 7). After

reacetylation, the GCMP was

diaf i ltered and concentrated

with a 30 kDa NMWCO UF

(Pel l icon 2 , Mi l l ipore) that

retain the GCMP. The GCMP

was then tested for concentra-

tion and purity.

T he G C M P conte nt w a s

determined by a color imet-

ric resorcinol-HCl method (1).

This method measured GCMP

monomer (sia l ic ac id) using

N-acetyl neuraminic acid as a

standard.The protein imputity

content was determined by the

Bradford method (2) using BSA

as a standard.

The nucleic acid impurity

content was determined by the

absorbance at 260 nm, assum-

ing an absorbance of 1 (1-cm

l ight path) for 50 µg/mL of

nucleic acid (3).

The white, waxy precipitate

was analyzed by Energy disper-

sive x-ray spectroscopy (EDXS),

micro-Fourier transform infra-

red (FT-IR) spectroscopy, proton

nuclear magnet ic resonance AL

L F

IGU

RE

S A

RE

CO

UR

TE

SY

OF

TH

E A

UT

HO

RS



Fermentation

1 2 3 4 5

6

7 8

Harvest

GCMP in

permeate

0.22µm

�ltration

300K UF

retentate

pool lots

Deacetylation/

Saponi�cation

high temp and

NaOH

50K UF

<50ºC

retentate

30K UF

retentatePrecipitation removed in

subsequent process steps

Precipitation

Acetylation

of GCMP

Precipitation

Figure 2: GCMP purification process.

Fatty acid1968 Lewis (6)

Average of 3 strains

1970 Moss (7)Range of 5

strains

2000 Rahman (8)Range of 10

strains

2010 Baxter C11 production strain

Analyzed by MIDI

c10:0 (decanoic) — trace – trace

c12:0 (lauric) 9 12–22 trace 4–17

c12:0 (-OH) – 20–35 – 2–9

c14:0 (myristic) 11 7–13 10–21 9–12

c14:1 – – trace trace

c14:0 (-OH) – 1–4 – 1

c15:0 (pentadecanoic) – 1–3 – –

c16:0 (palmitic) 33 11–20 43–54 20–38

c16:1 (palmitoleic) 26 15–21 22–31 35–37

c16:2 – – trace –

c17:0 (heptadecanoic) – trace 2 –

c18:0 (stearic) trace 1 – trace

c18:1 (oleic) 8 1–8 5–15 3–9

c19:0 (nonadecanoic) – 1–2 – –

c20:0 (arachidic) – 1–2 – –

c22:0 (behenic) – 1-2 – –

Unidentified 13 1–2 – <2

Table I: Fatty acid content (%) of Neisseria meningitidis.

(NMR), and liquid chromatogra-

phy-mass spectrometry (LC-MS)

to determine its composition.

LC-MS analysis of the precipitate

was performed on a Waters 2695

HPLC system with a C18 column

and a Waters Q-TOf API-US mass

spectrometer.

The fatty acids were identified

at Microbial ID (MIDI). Samples

were saponified in NaOH at ele-

vated temperature and methyl-

ated. The fatty acid methyl ester

was extracted in an organic sol-

vent prior to injection into the

gas chromatograph. Fatty acid

identification is based on their

retention times compared to a

library of standard (4).

Results and discussion

Eight successive purification lots

produced in a recent campaign

had an atypical appearance (see



Figure 3: Damage to the harvest pump mating surfaces.

Lot number Protein (%) Nucleic acid (%)

Process step Figure 2, Step 8Specification ≤ 5 ≤ 51 < 1 < 0.22 < 1 < 0.23 < 1 < 0.24 < 1 < 0.25 < 1 < 0.2

Table II: In-Process testing of 30K retentate for purity of the precipitation lots:

residual protein and nucleic acid

This amount of

precipitation had not

been noted in

past commercial

production.

Figure 1). These lots were cloudy

liquids with a white, waxy pre-

cipitate in the 50K retentate (see

Figure 2, step 6) and/or the 30K

retentate (see Figure 2, step 8).

The precipitate was identified as

mostly the sodium salt of palmitic

acid (C16:0). This amount of pre-

cipitation had not been noted in

past commercial production and

its appearance caused cessation

of manufacturing while this issue

was investigated.

It is believed that most of the

GCMP isolated from the fermenta-

tion is a lipidated molecule that is

able to aggregate either with itself

or with other macromolecules

such as lipopolysaccharide (5).

The aggregates are small enough

to pass through the 0.2 µm har-

vest filters but are retained by the

300 KDa UF (see Figure 2, step 4).

In the deacetylation, many

macromolecules are hydrolyzed

and the sodium salts of fatty acids

are generated. The most common

fatty acids in Neisseria are palmitic

and palmitoleic acids (see Table

I). The fatty acid analysis of our

working cell banks was consis-

tent with the literature (6-8) (see

Table I) and allowed the authors

to conclude that the source of the

precipitate was not exogenous but

was derived from the cells in our

fermentation.

During early development,

fatty acid characterization was

part of the product impurity

profile. Purified GCMP prior to

conjugation contained <0.5%

(w/w) fatty acids (palmitic, oleic,

and 3-OH myristic). After the

initial characterization, fatty

acids were not routinely ana-

lyzed, although residual protein

and nucleic acid levels were.

These data confirmed that most

fatty acids present were palmitic

acid, the most prevalent fatty

acid in Neisseria. Since the pre-

cipitate appeared in the 50K and

30K retentates, it was initially

thought that there were flaws in

the saponification and diafiltra-

tion steps. A significant portion

of the root cause investigation

examined the purification pro-

cess, but nothing unusual about

the deacetylation or saponifica-

tion was uncovered. The in-pro-

cess testing for residual protein

and nucleic acid gave acceptable

and typical results (see Table II).

Since no protein or nucleic acids

could be detected, and purifi-

cation of GCMP was achieved,

abnormalities in GCMP purifi-

cation were ruled out as a root

cause and the focus of the inves-

tigation shifted to the fermenta-

tion and harvest steps.

Many of the fermentation

inputs can affect cell metabolism

and as a consequence, GCMP

yield. Prior to the appearance of

precipitation, the authors had

seen increasing GCMP yields for

several months. It was hypothe-

sized that fermentation medium



It is reasonable

to conclude that

the damaged pump

affected the

integrity of the

cells by increasing

shear forces.

0

10

20

30

40

50

60

0 10 20 30 40 50 60

Fa

tty a

cid

/GC

MP

(%

w/w

)

Lot number index

Fermentation lots after implementingcorrective actions for PPT

Fermentation lots prior to observed PPT

Fermentation lots during observed PPT

Figure 4: Fatty acid/GCMP % (w/w) in fermentation lots.

Sample no.[Palmitic acid]

g/LCondition following

saponificationn

1 0.77

Precipitate formed

2 0.863 0.914 1.025 0.736 0.787 0.40

No precipitate8 0.359 0.3610 0.38

Table III: Concentration of palmitic acid entering the saponification step.

or operational parameters had

changed, increasing GCMP yield

in the form of lipidated-GCMP

and therefore overwhelming the

purification system. Several test

runs were made to alter the fer-

mentation process and reduce

the GCMP yield, but had no

apparent effect on the precipita-

tion problem.

The GCMP harvest was more

closely examined. This process

involved circulating the con-

tents of the fermentation ves-

sel through 0.2 µm hollow fiber

cartridges using a circumferen-

tial piston pump. This pump is

designed with moving part tol-

erances tighter than most sim-

ilarly-sized rotary lobe pumps.

Since the product is in the fil-

ter permeate, the fermentation

medium is continuously cir-

culated until the retained vol-

ume was low. We est imate

that each cell passes through

the harvest pump ~300 times.

Upon examination, the harvest

pump showed damage on mat-

ing surfaces in the lobes and

rotor housing (see Figure 3). This

damage had occurred when a

catastrophic event scored these

surfaces, rather than being the

result of normal wear and tear.

The pump continued to deliver

expected volumes and pres-

sures but the authors decided to

replace it with an identical pump.

The abnormal precipitation in

the downstream process immedi-

ately ceased.

The authors concluded that

the harvest pump, while opera-

tional, was damaged in such a

way that it was also acting as

a cell disruptor. Not only was

more lipidated-GCMP sheared

from the cell capsule, but the

cell membranes were disrupted,

forming fragments small enough

to pass though the 0.2 µm har-

vest filters (see Figure 2, steps 2

and 3), but large enough to be

retained in the initial GCMP

capture step (see Figure 2, step

4). Fatty acids in this mate-

rial would be saponified in the

deacetylation step (see Figure 2,

step 5). If the sodium salts were

soluble, they would be removed

from the system during diafil-

tration. If they were not solu-

ble, such as the sodium salt of

palmitic acid, they would pre-

cipitate during this process.

The abi l ity of the damaged

pump to mechanically extract

fatty acids from the cells must

have been increasing for several

months before a critical point

was reached where the amount

or type entering the downstream

process overwhelmed its ability

to remove it.



Fatty acid Sample

300K retentate 50K retentate 30K retentate

c12:0 31 1 not detectedc12:0 (OH) 19 trace not detectedc14:0 10 2 1c14:0 (3-OH) or c16:1 (iso) 2 3 2c16:0 11 80 88c16:1 24 5 2c18:0 not detected 2 2c18:1 2 6 5balance to 100% 1 1 0

Table IV: Fatty acid composition (%) at different stages of processing in lots that precipitated fatty acid.

Analysis of the 300K reten-

tates (see Figure 2, step 4) by GC

showed a general trend of lower

total fatty acid/GCMP ratios in

lots that did not generate precip-

itation downstream (see Figure

4). No specific level could pre-

dict if a lot would precipitate, but

30-40 % (w/w), could be used as

an alert level. This interpretation

was enhanced when the specific

concentration of palmitic acid

entering the deacetylation/sapon-

ification process was examined.

Deacetylations with palmitic acid

>0.7 g/l were likely to precipitate

(see Table III).

The 300K retentate was the

starting material for the saponi-

f icat ion react ion. There is a

mixture of saturated (palmitic,

myristic) and unsaturated (pal-

mitoleic, oleic) fatty acids in

this retentate (see Table IV).

The precipitate in both the 50K

and 30K retentates was com-

posed pr imar i ly of sod ium

pa lmitate . Th is cont rast i s

explained by the differences

in solubility of the fatty acids.

As reported by McBain et al9,

the sodium salts of unsaturated

fatty acids were more soluble

than the sodium salts of satu-

rated fatty acids. Another factor

that affects the fatty acid com-

position of the precipitate is the

length of the fatty acid hydro-

carbon chain: the longer the

hydrocarbon chain, the less sol-

uble the fatty acid was in water.

Sodium oleate and sodium lau-

rate are soluble in water at tem-

peratures under 45˚C. Sodium

myristate, palmitate, and stea-

rate have much lower solubility

at the same temperature. This

property may be the reason for

the oleic acid to be the more

abundant unsaturated fat t y

acid in the precipitate, despite

the fact that palmitoleic acid is

one of the most abundant fatty

acid in the 300K retentate.

Since an increase in fatty

acid concentration in the 300K

retentate was associated with

the use of a damaged pump, it

is reasonable to conclude that

the damaged pump af fected

the integr ity of the cells by

increasing shear forces due to

metal-to -metal contact. The

issue described here highlights

the importance of preventive

maintenance. The pump had

been in place for several years

and appeared to work properly,

in that it maintained normal

flow rates and pressures during

operation. Maintenance tech-

nicians performed preventive

maintenance at routine inter-

vals, focusing on electric cur-

rent demand and replac ing

hydraulic oil in the gear case.

The pump head was not rou-

tinely examined. Given what

was learned about the sensitiv-

ity of bacterial cells to shear,

t he aut hor s now v i s u a l ly

inspect the pump more often

and replace it at any sign of

damage.

AcknowledgmentsThe authors wish to thank Liqiong

Fang, PhD, research scientist, Kirk

Ashland, PhD, senior research sci-

entist, Frank Hua, PhD, research

scientist, and Catherine Quinn,

research associate, all at Baxter

Healthcare, Round Lake. IL, who

provided critical analyses in sup-

port of this study. BP

References 1. L. Svennerholm, Biochimica et

Biophysica Acta, 24, 604-11 (1957).

2. M. Bradford, Anal. Biochem. 72, 248–

54 (1976).

3. C.E. Frasch, Production and control of

Neisseria meningitidis vaccine, Adv

Biotechnol Processes. 13, 123–145

(1990).

4. Microbial Identification by Gas

Chromatographic Analysis of Fatty

Acid Methyl Esters (GC-FAME), MIDI

Technical note #101.

5. E.C. Gotschlich, et. al., J. Biol. Chem.

256, 8915–8921 (1981).

6. V.J. Lewis, R.E. Weaver, and D.G. Hollis,

J. Bacteriol. 96, 1–5 (1968).

7. C.W. Moss, et. al., J. Bacteriol. 104,

63–68 (1970).

8. M.M. Rahman, V.S.K. Kolli, C.M. Kahler,

G. Shih, D.S. Stephens, and R.W.

Carlson, Microbiol. 146, 1901–1911

(2000).

9. J.W. McBain, and W.C. Sierichs, J.

Amer. Oil Chemist’s Soc. 25, 221–225

(1948).

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