Advances in
Separation & Purification
April 2011
BioPharmINTERNATIONAL
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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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The Science & Business of Biopharmaceuticals
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.
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-
s6 Supplement to BioPharm International April 2011 www.biopharminternational.com
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
urt
esy A
uth
or/
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rto
riu
s
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s8 Supplement to BioPharm International April 2011 www.biopharminternational.com
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-
www.biopharminternational.com April 2011 Supplement to BioPharm International s9
Separation and Purification Challenges
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
s10 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Challenges
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
www.biopharminternational.com April 2011 Supplement to BioPharm International s11
Separation and Purification Challenges
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
s12 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Challenges
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
www.biopharminternational.com April 2011 Supplement to BioPharm International s13
Separation and Purification Challenges
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-
s14 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Challenges
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
www.biopharminternational.com April 2011 Supplement to BioPharm International s15
Separation and Purification Challenges
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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s16 Supplement to BioPharm International April 2011 www.biopharminternational.com
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-
www.biopharminternational.com April 2011 Supplement to BioPharm International s17
Separation and Purification Anion Exchange
* 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
s18 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Anion Exchange
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-
www.biopharminternational.com April 2011 Supplement to BioPharm International s19
Separation and Purification Anion Exchange
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.
s20 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Anion Exchange
* 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
www.biopharminternational.com April 2011 Supplement to BioPharm International s21
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.
s22 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Vaccines
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
www.biopharminternational.com April 2011 Supplement to BioPharm International s23
Separation and Purification Vaccines
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
s24 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Vaccines
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
www.biopharminternational.com April 2011 Supplement to BioPharm International s25
Separation and Purification Vaccines
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.
s26 Supplement to BioPharm International April 2011 www.biopharminternational.com
Separation and Purification Vaccines
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).
Macron™ Chemicals, formerly Mallinckrodt® Chemicals:
The new name for consistency in your application.
Formerly marketed under the Mallinckrodt® Chemicals brand name,
Macron™ Chemicals products have a long and successful history in demanding
laboratory and pharmaceutical markets. Today, the brand’s focus includes
products for everyday use in pharmaceutical and biopharmaceutical research
and manufacturing. Macron™ high purity solvents, acids, solutions, excipients,
salts and sugars will be produced in the same facilities using the same
manufacturing processes — and deliver the same high quality and purity
that you’ve come to expect.
Learn more at www.avantormaterials.com/macron-biopharm
Follow us @avantor_news
Mallinckrodt® is a trademark of Mallinckrodt Inc.Other trademarks are owned by Avantor Performance Materials, Inc. or its affi liates unless otherwise noted.© 2011 Avantor Performance Materials, Inc.
Now part of
performance has a new name.
Tosoh Bioscience now guarantees delivery* within 5 days or less to your facility of:
Toyopearl GigaCap S-650M
Toyopearl GigaCap Q-650M
Toyopearl GigaCap CM-650M
Your resin screening is complete and the decision is to go to GMP production. The pressure is on you. Sometimes weeks can be saved by using a supplier who maintains inventory for fast delivery.
If these Toyopearl® ion exchange resins are not delivered in 5 days or less, we will credit your invoice 5%. Toyopearl GigaCap resins are inventoried in our Grove City, Ohio warehouse and will ship the next business day after receipt of order.
TOSOH BIOSCIENCE LLC3604 Horizon Drive, Suite 100King of Prussia, PA 19406 Tel: 800-366-4875 email: [email protected]
Contact Tosoh Bioscience today or visit
our website to get more information about
Toyopearl GigaCap resin.
Tosoh Bioscience, Toyopearl and Toyopearl GigaCap are
registered trademarks of Tosoh Corporation.
I Need Delivery of my
Toyopearl GigaCap®
Resin NOW!
*This delivery guarantee applies to Toyopearl GigaCap products in 1L-15L quantities (in package sizes of 1L and 5L) delivered within the continental United States only. Other products and delivery locations excluded.