U.S. patent application number 12/389140 was filed with the patent office on 2009-08-27 for process for concentration of antibodies and therapeutic products thereof.
This patent application is currently assigned to Genentech, Inc.. Invention is credited to Charles M. WINTER.
Application Number | 20090214522 12/389140 |
Document ID | / |
Family ID | 35996499 |
Filed Date | 2009-08-27 |
United States Patent
Application |
20090214522 |
Kind Code |
A1 |
WINTER; Charles M. |
August 27, 2009 |
PROCESS FOR CONCENTRATION OF ANTIBODIES AND THERAPEUTIC PRODUCTS
THEREOF
Abstract
The present disclosure provides a process for concentrating
proteins including an ultrafiltering, a diafiltering, and a second
ultrafiltering sequence, at elevated temperatures, such as above
about 30.degree. C. The disclosure also includes a process for
preparing highly concentrated antibody compositions, and highly
concentrated antibody products.
Inventors: |
WINTER; Charles M.;
(Belmont, CA) |
Correspondence
Address: |
MORRISON & FOERSTER LLP
755 PAGE MILL RD
PALO ALTO
CA
94304-1018
US
|
Assignee: |
Genentech, Inc.
South San Francisco
CA
Novartis AG
Basel
|
Family ID: |
35996499 |
Appl. No.: |
12/389140 |
Filed: |
February 19, 2009 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
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11220362 |
Sep 6, 2005 |
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12389140 |
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60609092 |
Sep 9, 2004 |
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Current U.S.
Class: |
424/130.1 ;
530/414 |
Current CPC
Class: |
A61K 39/39591 20130101;
C07K 16/065 20130101; B01D 61/16 20130101; B01D 2311/16 20130101;
B01D 2311/04 20130101; C07K 2317/21 20130101; B01D 61/142 20130101;
C07K 16/4291 20130101; C07K 1/34 20130101; B01D 2311/04 20130101;
B01D 2311/10 20130101; B01D 2311/246 20130101 |
Class at
Publication: |
424/130.1 ;
530/414 |
International
Class: |
A61K 39/395 20060101
A61K039/395; C07K 1/34 20060101 C07K001/34 |
Claims
1. A process for preparing highly concentrated antibody
compositions comprising: a first ultrafiltering of a first antibody
preparation to provide a second antibody preparation; a
diafiltering the second antibody preparation to provide a
diafiltered intermediate antibody preparation; and a second
ultrafiltering of the diafiltered intermediate antibody preparation
to provide a third antibody preparation; wherein one or more of the
first ultrafiltering, second ultrafiltering, and the diafiltering
are accomplished at about 30.degree. C. to about 50.degree. C.
2. The process of claim 1 wherein one or more of the first
ultrafiltering, the second ultrafiltering, and the diafiltering are
accomplished at from about 35.degree. C. to about 50.degree. C.
3. The process of claim 1 wherein one or more of the first
ultrafiltering, the second ultrafiltering, and the diafiltering are
accomplished at from about 45.degree. C.
4. The process of claim 1 wherein one or more of the first
ultrafiltering, the second ultrafiltering, and the diafiltering are
accomplished at 45.degree. C. plus or minus 5.degree. C.
5. The process of claim 1 wherein the first antibody preparation
has an antibody concentration of from about 0.1 to about 10 grams
per liter.
6. The process of claim 1 wherein the first antibody preparation
has an antibody concentration of about 1 to about 5 grams per
liter.
7. The process of claim 1 wherein the second antibody preparation
has an antibody concentration of from about 10 to about 50 grams
per liter.
8. The process of claim 1 wherein the second antibody preparation
has an antibody concentration of about 20 to about 40 grams per
liter.
9. The process of claim 1 wherein the third antibody preparation
has an antibody concentration of from about 50 to about 250 grams
per liter.
10. The process of claim 1 wherein the third antibody preparation
has an antibody concentration of from about 100 to about 230 grams
per liter.
11. The process of claim 1 wherein the third antibody preparation
has an antibody concentration of from about 170 to about 200 grams
per liter.
12. The process of claim 1 wherein the diafiltered intermediate
antibody preparation and the third antibody preparation comprise
the ultra-filter retentate.
13. The process of claim 1 wherein the intermediate antibody
preparation has an antibody concentration of about 25 to about 35
grams per liter and the third antibody preparation has an antibody
concentration of from about 170 to about 200 grams per liter.
14. The process of claim 1 wherein the antibody preparation
comprises anti-IgE antibodies.
15. The process of claim 1 wherein the process is accomplished in
from about 1 to 10 hours.
16. The process of claim 1 wherein the process is accomplished in
from about 2 to 5 hours.
17. The process of claim 1 wherein the process is accomplished in
about 3 hours.
18. The process of claim 1 wherein the first and the second
ultrafiltering are accomplished with an ultra-filter membrane
having a nominal pore size of about 5 to about 50 kilo Daltons.
19. The process of claim 1 wherein the first and the second
ultrafiltering are accomplished with an ultra-filter membrane
having a nominal pore size of about 10 to about 30 kilo
Daltons.
20. The process of claim 1 wherein the first antibody preparation
contains an antibody having an apparent molecular weight of about
100 to about 200 kilo Daltons.
21. The process of claim 1 wherein the first antibody preparation
contains an antibody having an apparent molecular weight of about
150 kilo Daltons.
22. The process of claim 1 wherein the first ultrafiltering
concentrates the first antibody preparation to provide the second
antibody preparation having an antibody concentration of about 30
grams per liter and the second ultrafiltering concentrates the
intermediate antibody preparation to provide the third antibody
preparation having an antibody concentration of about 170 to about
200 grams per liter.
23. The process of claim 1 wherein the first ultrafiltering and the
second ultrafiltering are accomplished with the same ultra-filter
membrane.
24. The process of claim 1 wherein the first ultrafiltering and the
second ultrafiltering are accomplished with a regenerated cellulose
composite ultra-filter membrane.
25. The process of claim 1 wherein the diafiltering accomplishes a
buffer exchange at constant volume, constant concentration, or
both.
26. The process of claim 1 wherein the diafiltering accomplishes a
buffer exchange of from about 5 to about 15 fold volumes.
27. The process of claim 1 wherein the diafiltering accomplishes a
buffer exchange of from about 8 fold volumes.
28. The process of claim 1 wherein the diafiltering exchanges a
first buffer for a second buffer.
29. The process of claim 28 wherein the first buffer comprises a
mixture of aqueous sodium chloride and a TRIS buffer, and the
second buffer comprises a mixture of aqueous histidine chloride and
arginine chloride.
30. The process of claim 1 wherein the first ultrafiltering, the
second ultrafiltering, and the diafiltering are accomplished with
tangential flow filtration across an ultra-filter membrane.
31. The process of claim 1 wherein the first ultrafiltering, the
second ultrafiltering, and the diafiltering are accomplished with
tangential flow filtration across the same ultra-filter
membrane.
32. The process of claim 1 wherein the yield of the third antibody
preparation is greater than about 70 weight % based on the weight
of antibodies in the first antibody preparation.
33. The process of claim 32 wherein the yield of the third antibody
preparation is from about 80 to about 100 weight % based on the
weight of antibodies in the first antibody preparation.
34. The process of claim 1 wherein the first ultrafiltering has a
recirculation rate of from about 0.5 L/min/ft.sup.2 to about 5
L/min/ft.sup.2.
35. The process of claim 1 wherein the ultrafiltering and
diafiltering are accomplished at a transmembrane pressure of from
about 10 to about 50 p.s.i.
36. The process of claim 1 wherein there is provided an antibody
concentrate with a detectable bio-burden of less than about 100
CFU/mL.
37. A process for concentrating proteins comprising: a first
ultrafiltering of a first protein mixture to provide a second
protein mixture; a diafiltering of the second protein mixture to
provide a diafiltered protein mixture; and a second ultrafiltering
of the diafiltered protein mixture to provide a third protein
mixture, wherein one or more of the first ultrafiltering, the
diafiltering, and the second ultrafiltering are accomplished at
from about 30.degree. C. to about 50.degree. C.
Description
RELATED APPLICATIONS
[0001] This application claims the priority benefit of U.S.
Provisional Application Ser. No. 60/609,092 filed Sep. 9, 2004,
which is incorporated herein by reference in its entirety.
BACKGROUND
[0002] Methods for isolating, purifying, and concentrating
biological materials are known and include, for example,
chromatography, ultrafiltration, and lyophilization, see generally,
R. Hatti-Kaul et al., "Downstream Processing in Biotechnology," in
Basic Biotechnology, Chap. 9, pages 187-211, 2nd ed., Cambridge
University Press (2001). Processes for making concentrated
monoclonal antibody preparations for administration to humans are
known, see for example, U.S. Pat. No. 6,252,055, which uses
ultrafiltration and which re-circulates the resulting filtrate.
[0003] Some challenges associated with available antibody
concentration methods include, for example, low fluxes, long
process times, large membrane areas, mechanical recovery yield and
losses, operator-intensive intervention or handling, low mass
transfer rates, energy inefficiencies, and hydraulic pressure
limits on concentration equipment. These and other challenges can
contribute to a high total cost of manufacture and ultimately
higher costs to therapeutic drug consumers.
[0004] There is a need for improved processes for preparing highly
concentrated protein formulations, such as liquid antibody
preparations and therapeutic products thereof.
SUMMARY
[0005] In general terms, the present disclosure generally relates
to processes for concentrating proteins, such as processes for
concentrating an antibody preparation, pharmaceutical formulations
containing such a preparation, and there use in human therapy or
animal therapy.
[0006] In embodiments the present disclosure provides processes for
preparing highly concentrated proteins, such as antibody
preparations; and therapeutic products prepared by the process,
such as therapeutic antibody products. Accordingly, the present
disclosure provides, a process for concentrating proteins
comprising: a first ultrafiltering of a first antibody preparation
to provide a second antibody preparation; a diafiltering the second
antibody preparation to provide a diafiltered intermediate antibody
preparation; and a second ultrafiltering of the diafiltered
intermediate antibody preparation to provide a third antibody
preparation, wherein one or more of the first ultrafiltering, the
second ultrafiltering, and the diafiltering are accomplished at
elevated temperatures, for example, from about 30.degree. C. to
about 50.degree. C.
[0007] The present disclosure also provides, in embodiments, a
process for concentrating proteins comprising: a first
ultrafiltering of a first protein mixture to provide a second
protein mixture; a diafiltering the second protein mixture to
provide a diafiltered protein mixture; and a second ultrafiltering
of the diafiltered protein mixture to provide a third protein
mixture, wherein one or more of the first ultrafiltering, the
diafiltering, and the second ultrafiltering are accomplished at,
for example, about 45.degree. C.
[0008] The present disclosure also provides, in embodiments, a
highly concentrated antibody composition prepared by the above
processes.
BRIEF DESCRIPTION OF THE DRAWINGS
[0009] FIG. 1 illustrates an apparatus for accomplishing the
preparative process, in embodiments of the present disclosure.
[0010] FIGS. 2 through 17 illustrate various observed or measured
process values over various phases or mode of the process, in
embodiments of the present disclosure.
[0011] FIGS. 18 and 19 illustrate the effect of elevated
temperature on product quality, in embodiments of the present
disclosure.
[0012] FIGS. 20 and 21 illustrate the effect of elevated
temperature on bioburden control, in embodiments of the present
disclosure.
[0013] FIG. 22 illustrates the effect of elevated temperature on
process flux and process time, in embodiments of the present
disclosure.
[0014] FIGS. 23 through 25 illustrates various observed or measured
process values over various phases or mode of the scaled-up
process, in embodiments of the present disclosure.
DETAILED DESCRIPTION
[0015] Various embodiments of the present disclosure will be
described in detail with reference to drawings, if any. Reference
to various embodiments does not limit the scope of the invention,
which is limited only by the scope of the claims attached hereto.
Additionally, any examples set forth in this specification are not
intended to be limiting and merely set forth some of the many
possible embodiments for the claimed invention.
[0016] The following are used, unless otherwise described:
[0017] "Ultrafiltering," "ultrafiltration," "ultrafiltered," "UF,"
and like terms refer to, for example, using synthetic
semi-permeable membranes, with appropriate physical and chemical
properties, to discriminate between molecules in the mixture,
primarily on the basis of molecular size and shape, and accomplish
separation of different molecules or accomplish concentration of
like molecules.
[0018] "Diafiltering," "diafiltration," "diafiltered,"
"diafiltrating," "DF," and like terms refer to, for example, using
an ultrafiltration membrane to remove, replace, or lower the
concentration of salts or solvents from solutions or mixtures
containing proteins, peptides, nucleic acids, or other
biomolecules.
[0019] "Transmembrane pressure" or "TMP" refers to the average
applied pressure from the feed to the filtrate side of the membrane
calculated as TMP [bar]=[(P.sub.F+P.sub.R)/2]-P.sub.f. where
P.sub.F is the feed pressure, P.sub.R is the retentate pressure,
and P.sub.f is the filtrate pressure.
[0020] "Tangential flow filtration," "cross flow filtration,"
"TFF," and like terms refer to a mode of filtration in which the
solute-containing solution passes tangentially across the UF
membrane and lower molecular weigh salts or solutes are passed
through by applying pressure.
[0021] "Antibody" is used in the broadest sense and specifically
covers intact monoclonal antibodies, polyclonal antibodies,
multispecific antibodies (e.g., bispecific antibodies) formed from
at least two intact antibodies, and antibody fragments, so long as
they exhibit the desired biological activity. An antibody is a
protein generated by the immune system that is capable of
recognizing and binding to a specific antigen. Described in terms
of its structure, an antibody is a Y-shaped protein consisting of
four amino acid chains, two heavy and two light. In a simplified
model sufficient for this appeal, each antibody has primarily two
regions: a variable region and a constant region. The variable
region, located on the ends of the arms of the Y, binds to and
interacts with the target antigen. This variable region includes a
complementary determining region (CDR) that recognizes and binds to
a specific binding site on a particular antigen. The constant
region, located on the tail of the Y, is recognized by and
interacts with the immune system (Janeway, C., Travers, P.,
Walport, M., Shlomchik (2001) Immuno Biology, 5th Ed., Garland
Publishing, New York). A target antigen generally has numerous
binding sites, also called epitopes, recognized by CDRs on multiple
antibodies. Each antibody that specifically binds to a different
epitope has a different structure. Thus, one antigen may have more
than one corresponding antibody.
[0022] The basic 4-chain antibody unit is a heterotetrameric
glycoprotein composed of two identical light (L) chains and two
identical heavy (H) chains (an IgM antibody consists of 5 of the
basic heterotetramer unit along with an additional polypeptide
called J chain, and therefore contain 10 antigen binding sites,
while secreted IgA antibodies can polymerize to form polyvalent
assemblages comprising 2-5 of the basic 4-chain units along with J
chain). In the case of IgGs, the 4-chain unit is generally about
150,000 daltons. Each L chain is linked to an H chain by one
covalent disulfide bond, while the two H chains are linked to each
other by one or more disulfide bonds depending on the H chain
isotype. Each H and L chain also has regularly spaced intrachain
disulfide bridges. Each H chain has at the N-terminus, a variable
domain (V.sub.H) followed by three constant domains (C.sub.H) for
each of the .alpha. and .gamma. chains and four C.sub.H domains for
.mu. and .epsilon. isotypes. Each L chain has at the N-terminus, a
variable domain (V.sub.L) followed by a constant domain (C.sub.L)
at its other end. The V.sub.L is aligned with the V.sub.H and the
C.sub.L is aligned with the first constant domain of the heavy
chain (C.sub.H1). Particular amino acid residues are believed to
form an interface between the light chain and heavy chain variable
domains. The pairing of a V.sub.H and V.sub.L together forms a
single antigen-binding site. For the structure and properties of
the different classes of antibodies, see e.g., Basic and Clinical
Immunology, 8th edition, D. Stites, A. Terr and T. Parslow (eds.),
Appleton & Lange, Norwalk, Conn., 1994, page 71 and Chapter
6.
[0023] The L chain from any vertebrate species can be assigned to
one of two clearly distinct types, called kappa and lambda, based
on the amino acid sequences of their constant domains. Depending on
the amino acid sequence of the constant domain of their heavy
chains (C.sub.H), immunoglobulins can be assigned to different
classes or isotypes. There are five classes of immunoglobulins:
IgA, IgD, IgE, IgG, and IgM, having heavy chains designated
.alpha., .delta., .epsilon., .epsilon., and .mu., respectively. The
.gamma. and .alpha. classes are further divided into subclasses on
the basis of relatively minor differences in C.sub.H sequence and
function, e.g., humans express the following subclasses: IgG1,
IgG2, IgG3, IgG4, IgA1, and IgA2.
[0024] The term "variable" refers to the fact that certain segments
of the variable domains differ extensively in sequence among
antibodies. The V domain mediates antigen binding and define
specificity of a particular antibody for its particular antigen.
However, the variability is not evenly distributed across the
approximately 110-amino acid span of the variable domains. Instead,
the V regions consist of relatively invariant stretches called
framework regions (FRs) of 15-30 amino acids separated by shorter
regions of extreme variability called "hypervariable regions" that
are each 9-12 amino acids long. The variable domains of native
heavy and light chains each comprise four FRs, largely adopting a
.beta.-sheet configuration, connected by three hypervariable
regions, which form loops connecting, and in some cases forming
part of, the .beta.-sheet structure. The hypervariable regions in
each chain are held together in close proximity by the FRs and,
with the hypervariable regions from the other chain, contribute to
the formation of the antigen-binding site of antibodies (see Kabat
et al., in Sequences of Proteins of Immunological Interest, 5th Ed.
Public Health Service, National Institutes of Health, Bethesda, Md.
(1991)). The constant domains are not involved directly in binding
an antibody to an antigen, but exhibit various effector functions,
such as participation of the antibody in antibody dependent
cellular cytotoxicity (ADCC).
[0025] The term "hypervariable region" when used herein refers to
the amino acid residues of an antibody which are responsible for
antigen-binding. The hypervariable region generally comprises amino
acid residues from a "complementarity determining region" or "CDR"
(e.g., around about Kabat residues 24-34 (L1), 50-56 (L2) and 89-97
(L3) in the V.sub.L, and around about Kabat residues 31-35B (H1),
50-65 (H2) and 95-102 (H3) in the V.sub.H (see Kabat et al., supra)
and/or those residues from a "hypervariable loop" (e.g., around
about Chothia residues 26-32 (L1), 50-52 (L2) and 91-96 (L3) in the
V.sub.L, and 26-32 (H1), 52A-55 (H2) and 96-101 (H3) in the V.sub.H
(Chothia and Lesk, J. Mol. Biol, 196:901-917 (1987)).
[0026] The term "monoclonal antibody" as used herein refers to an
antibody from a population of substantially homogeneous antibodies,
i.e., the individual antibodies comprising the population are
identical and/or bind the same epitope(s), except for possible
variants that may arise during production of the monoclonal
antibody, such variants generally being present in minor amounts.
Such monoclonal antibody typically includes an antibody comprising
a polypeptide sequence that binds a target, wherein the
target-binding polypeptide sequence was obtained by a process that
includes the selection of a single target binding polypeptide
sequence from a plurality of polypeptide sequences. For example,
the selection process can be the selection of a unique clone from a
plurality of clones, such as a pool of hybridoma clones, phage
clones or recombinant DNA clones. It should be understood that the
selected target binding sequence can be further altered, for
example, to improve affinity for the target, to humanize the target
binding sequence, to improve its production in cell culture, to
reduce its immunogenicity in vivo, to create a multispecific
antibody, etc., and that an antibody comprising the altered target
binding sequence is also a monoclonal antibody of this invention.
In contrast to polyclonal antibody preparations which typically
include different antibodies directed against different
determinants (epitopes), each monoclonal antibody of a monoclonal
antibody preparation is directed against a single determinant on an
antigen. In addition to their specificity, the monoclonal antibody
preparations are advantageous in that they are typically
uncontaminated by other immunoglobulins. The modifier "monoclonal"
indicates the character of the antibody as being obtained from a
substantially homogeneous population of antibodies, and is not to
be construed as requiring production of the antibody by any
particular method. For example, the monoclonal antibodies to be
used in accordance with the present invention may be made by a
variety of techniques, including, for example, the hybridoma method
(e.g., Kohler et al., Nature, 256:495 (1975); Harlow et al.,
Antibodies: A Laboratory Manual, (Cold Spring Harbor Laboratory
Press, 2nd ed. 1988); Hammerling et al., in: Monoclonal Antibodies
and T-Cell Hybridomas, 563-681, (Elsevier, N.Y., 1981)),
recombinant DNA methods (see, e.g., U.S. Pat. No. 4,816,567), phage
display technologies (see, e.g., Clackson et al., Nature,
352:624-628 (1991); Marks et al., J. Mol. Biol., 222:581-597
(1991); Sidhu et al., J. Mol. Biol. 338(2):299-310 (2004); Lee et
al., J. Mol. Biol. 340(5):1073-1093 (2004); Fellouse, Proc. Nat.
Acad. Sci. USA 101(34):12467-12472 (2004); and Lee et al., J.
Immunol. Methods 284(1-2):119-132 (2004), and technologies for
producing human or human-like antibodies in animals that have parts
or all of the human immunoglobulin loci or genes encoding human
immunoglobulin sequences (see, e.g., WO 1998/24893; WO 1996/34096;
WO 1996/33735; WO 1991/10741; Jakobovits, et al., Proc. Natl. Acad.
Sci. USA, 90:2551 (1993); Jakobovits, et al., Nature, 362:255-258
(1993); Bruggemann, et al., Year in Immuno., 7:33 (1993); U.S. Pat.
Nos. 5,545,806; 5,569,825; 5,591,669 (all to GenPharm); 5,545,807;
WO 1997/17852; U.S. Pat. Nos. 5,545,807; 5,545,806; 5,569,825;
5,625,126; 5,633,425; and 5,661,016; Marks, et al., Bio/Technology,
10: 779-783 (1992); Lonberg, et al., Nature 368: 856-859 (1994);
Morrison, Nature, 368: 812-813 (1994); Fishwild, et al., Nature
Biotechnology, 14: 845-851 (1996); Neuberger, Nature Biotechnology,
14: 826 (1996); and Lonberg and Huszar, Intern. Rev. Immunol., 13:
65-93 (1995).
[0027] "Chimeric" antibodies (immunoglobulins) have a portion of
the heavy and/or light chain identical with or homologous to
corresponding sequences in antibodies derived from a particular
species or belonging to a particular antibody class or subclass,
while the remainder of the chain(s) is identical with or homologous
to corresponding sequences in antibodies derived from another
species or belonging to another antibody class or subclass, as well
as fragments of such antibodies, so long as they exhibit the
desired biological activity (U.S. Pat. No. 4,816,567; and Morrison,
et al., Proc. Natl. Acad. Sci. USA 81:6851-6855 (1984)). Humanized
antibody as used herein is a subset of chimeric antibodies.
[0028] "Humanized" forms of non-human (e.g., murine) antibodies are
chimeric antibodies which contain minimal sequence derived from
non-human immunoglobulin. For the most part, humanized antibodies
are human immunoglobulins (recipient or acceptor antibody) in which
hypervariable region residues of the recipient are replaced by
hypervariable region residues from a non-human species (donor
antibody) such as mouse, rat, rabbit or nonhuman primate having the
desired specificity, affinity, and capacity. In some instances, Fv
framework region (FR) residues of the human immunoglobulin are
replaced by corresponding non-human residues. Furthermore,
humanized antibodies may comprise residues which are not found in
the recipient antibody or in the donor antibody. These
modifications are made to further refine antibody performance such
as binding affinity. Generally, the humanized antibody will
comprise substantially all of at least one, and typically two,
variable domains, in which all or substantially all of the
hypervariable loops correspond to those of a non-human
immunoglobulin and all or substantially all of the FR regions are
those of a human immunoglobulin sequence although the FR regions
may include one or more amino acid substitutions that improve
binding affinity. The number of these amino acid substitutions in
the FR are typically no more than 6 in the H chain, and in the L
chain, no more than 3. The humanized antibody optionally also will
comprise at least a portion of an immunoglobulin constant region
(Fc), typically that of a human immunoglobulin. For further
details, see Jones, et al., Nature 321:522-525 (1986); Reichmann,
et al., Nature 332:323-329 (1988); and Presta, Curr. Op. Struct.
Biol. 2:593-596 (1992).
[0029] "Antibody fragments" comprise a portion of an intact
antibody, preferably the antigen binding or variable region of the
intact antibody. Examples of antibody fragments include Fab, Fab',
F(ab').sub.2, and Fv fragments; diabodies; linear antibodies (see
U.S. Pat. No. 5,641,870, Example 2; Zapata, et al., Protein Eng.,
8(10): 1057-1062 (1995)); single-chain antibody molecules; and
multispecific antibodies formed from antibody fragments.
[0030] Papain digestion of antibodies produces two identical
antigen-binding fragments, called "Fab" fragments, and a residual
"Fc" fragment, a designation reflecting the ability to crystallize
readily. The Fab fragment consists of an entire L chain along with
the variable region domain of the H chain (VH), and the first
constant domain of one heavy chain (C.sub.H1). Each Fab fragment is
monovalent with respect to antigen binding, i.e., it has a single
antigen-binding site. Pepsin treatment of an antibody yields a
single large F(ab').sub.2 fragment which roughly corresponds to two
disulfide linked Fab fragments having divalent antigen-binding
activity and is still capable of cross-linking antigen. Fab'
fragments differ from Fab fragments by having additional few
residues at the carboxy terminus of the C.sub.H1 domain including
one or more cysteines from the antibody hinge region. Fab'-SH is
the designation herein for Fab' in which the cysteine residue(s) of
the constant domains bear a free thiol group. F(ab').sub.2 antibody
fragments originally were produced as pairs of Fab' fragments which
have hinge cysteines between them. Other chemical couplings of
antibody fragments are also known.
[0031] The Fc fragment comprises the carboxy-terminal portions of
both H chains held together by disulfides. The effector functions
of antibodies are determined by sequences in the Fc region, which
region is also the part recognized by Fc receptors (FcR) found on
certain types of cells.
[0032] "Fv" is the minimum antibody fragment which contains a
complete antigen-recognition and -binding site. This fragment
consists of a dimer of one heavy- and one light-chain variable
region domain in tight, non-covalent association. From the folding
of these two domains emanate six hypervariable loops (3 loops each
from the H and L chain) that contribute the amino acid residues for
antigen binding and confer antigen binding specificity to the
antibody. However, even a single variable domain (or half of an Fv
comprising only three CDRs specific for an antigen) has the ability
to recognize and bind antigen, although at a lower affinity than
the entire binding site.
[0033] "Single-chain Fv" also abbreviated as "sFv" or "scFv" are
antibody fragments that comprise the V.sub.H and V.sub.L antibody
domains connected into a single polypeptide chain. Preferably, the
sFv polypeptide further comprises a polypeptide linker between the
V.sub.H and V.sub.L domains which enables the sFv to form the
desired structure for antigen binding. For a review of sFv, see
Pluckthun in The Pharmacology of Monoclonal Antibodies, vol. 113,
Rosenburg and Moore eds., Springer-Verlag, New York, pp. 269-315
(1994).
[0034] "About" modifying, for example, the quantity of an
ingredient in the compositions, concentration of an active, buffer
volumes, diavolumes, pore size, apparent molecular, molecular
weight cut-off, process temperature, process time, yields, flow
rates, pressures, bio-burdens, and like values, and ranges thereof,
employed in the methods of the invention, refers to variation in
the numerical quantity that can occur, for example, through typical
measuring and handling procedures used for making concentrates or
use solutions; through inadvertent error in these procedures;
through differences in the manufacture, source, or purity of the
ingredients employed to make the compositions or carry out the
methods; and like considerations. The term "about" also encompasses
amounts that differ due to aging of a composition with a particular
initial concentration or mixture. The term "about" also encompasses
amounts that differ due to mixing or processing a composition with
a particular initial concentration or mixture. Whether or not
modified by the term "about" the claims include equivalents to the
quantities.
[0035] "Consisting essentially of" refers to a process of obtaining
a concentrated protein composition or antibody composition that
includes the steps and ingredients listed in the claim, plus other
steps and ingredients that do not materially affect the basic and
novel properties of the composition, such as a multiplicity of
steps or buffer media. Ingredients that materially affect the basic
properties of the composition and method of the present disclosure
impart undesirable characteristics including, for example,
bio-burden, such as the undesirable toxicity or irritability
associated with contaminants.
[0036] The indefinite article "a" or "an" and its corresponding
definite article "the" as used herein is understood to mean at
least one, or one or more, unless specified otherwise.
[0037] The present disclosure provides, in embodiments, the
abovementioned processes and the concentrated antibody products
thereof.
[0038] In embodiments of the present disclosure, the preparative
processes and products thereof can be used in preparing highly
concentrated antibody preparations and similar preparations, such
as purifying and concentrating proteins or like substances from
natural or synthetic sources, and which products can be useful for
treating pathological conditions, such as asthma, cancer,
psoriasis, inhibiting angiogenesis, and like pathological
conditions.
[0039] In embodiments of the above-mentioned process for preparing
highly concentrated antibody compositions of the disclosure, the
following further exemplifies how to make and use the preparative
processes and products of the disclosure.
[0040] In embodiments of the present disclosure, there is provided
a process for preparing highly concentrated antibody compositions,
for example, according to accomplishing the following steps in the
order recited, comprising:
[0041] a first ultrafiltering of a first antibody preparation,
having a concentration of, for example, about 0.1 to about 10 grams
per liter (g/L), to provide an second antibody preparation as the
retentate, having a greater antibody concentration of, for example,
about 10 to about 50 grams per liter;
[0042] a diafiltering of the resulting second antibody preparation
to provide a diafiltered intermediate antibody preparation as the
retentate, having about the same concentration as the resulting
second antibody preparation retentate, that is, diafiltering to
accomplish a buffer exchange at constant volume; and
[0043] a second ultrafiltering of the diafiltered intermediate
antibody preparation to provide a third antibody preparation as the
retentate, having a greater antibody concentration of, for example,
about 150 to about 200 grams per liter.
[0044] The preparative processes of the disclosure can further
comprise an optional product recovery step or steps, for example,
and as disclosed and illustrated herein.
[0045] In embodiments of the above-mentioned process of the
disclosure, one or more of the first ultrafiltering, the
diafiltering, and the second ultrafiltering, can be accomplished
at, for example, from about 30.degree. C. to about 70.degree. C. In
embodiments, these steps can also be accomplished at, for example,
from about 30.degree. C. to about 50.degree. C. In embodiments,
these steps can also be accomplished at, for example, from about
35.degree. C. to about 50.degree. C. In embodiments, these steps
can also be accomplished at, for example, about 45.degree. C., such
as from about 45.degree. C. plus or minus 5.degree. C. Depending
upon the type of antibody preparation, for processes accomplished
at temperatures above about 70.degree. C., the preparation may show
signs of deterioration, such as denaturation, agglomeration, and
like phenomena. For processes accomplished at temperatures below
from about 30 to about 35.degree. C., the flow rates are typically
undesirably low and process times are undesirably long, making the
process at lower temperatures less attractive for efficient
commercial production.
[0046] In embodiments, the first antibody preparation can have an
antibody concentration of, for example, from about 0.1 to about 100
grams per liter (g/L). The antibody concentration is, for example,
a common concentration typically available from other preliminary
protein or antibody purification steps or methods, such as,
centrifugation, filtration, chromatography, and like procedures.
The resulting second antibody preparation obtainable from the first
ultrafiltering can have an antibody concentration of, for example,
from about 10 to about 50 grams per liter, and for example, about
20 to about 40 grams per liter, such as 30 grams per liter. A range
for the antibody concentration of the intermediate antibody
preparation can depend upon, for example, a balance of factors,
such as sample volume and sample flux achievable with a particular
buffer containing the second antibody preparation. The intermediate
antibody preparation can have an antibody concentration of, for
example, about 25 to about 35 grams per liter and the third
antibody preparation can have an antibody concentration of, for
example, from about 170 to about 200 grams per liter. The third
antibody preparation, in embodiments, can have an antibody
concentration of, for example, from about 50 to about 250 grams per
liter, such as from of about 100 to about 230 grams per liter, and
from about 170 to about 200 grams per liter, such as 185 grams per
liter.
[0047] It will be apparent to one skilled in the art, upon
comprehending the present disclosure, that the intermediate
antibody preparation and third antibody preparation comprise the
same ultra-filtered retentate except for, for example, differences
in the antibody concentration resulting from the first and second
ultrafiltering concentration, and differences in the suspending
buffer media resulting from the diafiltration buffer exchange.
Thus, there is little, if any, compositional change, such as
degradation, of the target protein or antibody product, in
embodiments of the present disclosure.
[0048] Conventional ultrafiltration concentration methods can have
generally greater time and lesser through-put inefficiencies having
considerably longer process times such as several days to several
weeks, process considerably smaller volumes, or both.
[0049] In embodiments, the protein concentration process of the
disclosure can be accomplished in, for example, from about 1 to 10
hours, preferably in from about 2 to 5 hours, and more preferably
in about 3 hours. The preferences favor higher flux through-put and
smaller membrane areas.
[0050] In embodiments, the first ultrafiltering can be
accomplished, for example, in about 35 percent of the total process
time. Thus, for example, in a concentration and purification
process of the disclosure with about 3 hours total process time,
the first ultrafiltering can be accomplished in about 45 minutes.
In embodiments, the second ultrafiltering can be accomplished, for
example, in about 15 percent of the total process time. Thus, for
example, in a process of the disclosure with about 3 hours total
process time, the second ultrafiltering can be accomplished in
about 15 minutes. The diafiltering can be accomplished, for
example, in about 50 percent of the total process time. Thus, for
example, in a process of the disclosure with about 3 hours total
process time, the diafiltering can be accomplished in from about 90
to about 120 minutes.
[0051] In embodiments, the first ultrafiltering and the second
ultrafiltering can be accomplished, for example, with an
ultra-filter membrane having a nominal pore size, or molecular
weight cut-off, of about 5 to about 50 kiloDaltons. Another
suitable nominal pore size is, for example, from about 10 to about
40 kiloDaltons. Yet another suitable nominal pore size, or
molecular weight cut-off, is about 30 kiloDaltons.
[0052] In embodiments, the first antibody preparation can contain,
for example, an antibody having an apparent molecular weight of,
for example, about 100 to about 200 kiloDaltons. In other
embodiments, the first antibody preparation can contain an antibody
having an apparent molecular weight of, for example, about 150
kiloDaltons, such as when the antibody preparation comprises
anti-IgE antibodies or IgE, see for example, U.S. Pat. No.
6,172,213 assigned to Genentech, Inc.
[0053] Other antibodies suitable for use in the present disclosure
include cancer treating antibodies, see generally, for example:
PCT/US02/19592; PCT/US01/20118; PCT/US01/25464; PCT/US01/26626;
PCT/US02/28859; PCT/US02/41798; PCT/US02/12206; PCT/US03/11148;
PCT/US02/12619; and PCT/US02/33050. Still other antibodies suitable
for use in the present disclosure include an anti-CD20 antibody and
like antibodies including human, non-human, murine, hybrid, and
chimeric forms. See for example U.S. Pat. No. 6,582,959 (VEGF) and
U.S. Patent Application No. 2002/0122797 A1 (human VEGF).
[0054] In embodiments, antibodies included within the scope of the
disclosure include hybrid and recombinant antibodies (e.g.,
"humanized" and "human" antibodies) regardless of species of origin
or immunoglobulin class or subclass designation, as well as
antibody fragments (for example, Fab, F(ab').sub.2, and F.sub.v).
See U.S. Pat. No. 4,816,567; Mage and Lamoyi, in Monoclonal
Antibody Production Techniques and Applications, 79-97, Marcel
Dekker, Inc., New York, (1987).
[0055] Monoclonal antibodies may also be used and can be isolated
from phage antibody libraries using the techniques described in
Clackson et al (1991) Nature, 352:624-628 and Marks, et al. (1991)
J. Mol. Biol., 222:581-597, for example. Monoclonal antibodies
include "chimeric" antibodies in which a portion of the heavy
and/or light chain is identical with or homologous to corresponding
sequences in antibodies derived from a particular species or
belonging to a particular antibody class or subclass, while the
remainder of the chain(s) is identical with or homologous to
corresponding sequences in antibodies derived from another species
or belonging to another antibody class or subclass, as well as
fragments of such antibodies, so long as they exhibit the desired
biological activity (U.S. Pat. No. 4,816,567; and Morrison, et al.
(1984) Proc. Natl. Acad. Sci. USA, 81:6851-6855). Chimeric
antibodies can include "primatized" antibodies comprising variable
domain antigen-binding sequences derived from a non-human primate
(e.g., Old World Monkey, Ape etc) and human constant region
sequences.
[0056] Monoclonal antibodies are highly specific, being directed
against a single antigenic site. Furthermore, in contrast to
polyclonal antibody preparations that include different antibodies
directed against different determinants (epitopes), each monoclonal
antibody is directed against a single determinant on the antigen.
In addition to their specificity, the monoclonal antibodies are
advantageous in that they may be synthesized uncontaminated by
other antibodies. Thus, the modifier "monoclonal" indicates the
character of the antibody as being obtained from such a
substantially homogeneous population of antibodies, i.e., the
individual antibodies comprising the population are identical
except for possible naturally occurring mutations that may be
present in minor amounts, and is not to be construed as requiring
production of the antibody by any particular method. For example,
the monoclonal antibodies for use in the disclosure may be made
using the hybridoma method first described by Kohler &
Milstein, Nature, 256:495 (1975), or may be made by recombinant DNA
methods. Other known methods of antibody production are described,
for example, in Goding, Monoclonal Antibodies: Principles and
Practice, 59-103, Academic Press (1986); Kozbor, J. Immunol.,
133:3001 (1984). Brodeur, et al., Monoclonal Antibody Production
Techniques and Applications, 51-63, Marcel Dekker, Inc., New York
(1987).
[0057] Various methods have been employed to produce monoclonal
antibodies (MAbs). Hybridoma technology, which refers to a cloned
cell line that produces a single type of antibody, uses the cells
of various species, including mice (murine), hamsters, rats, and
humans. Another method to prepare MAbs uses genetic engineering
including recombinant DNA techniques. Monoclonal antibodies made
from these techniques include, among others, chimeric antibodies
and humanized antibodies. A chimeric antibody combines DNA encoding
regions from more than one type of species. For example, a chimeric
antibody may derive the variable region from a mouse and the
constant region from a human. A humanized antibody comes
predominantly from a human, even though it contains nonhuman
portions. Like a chimeric antibody, a humanized antibody may
contain a completely human constant region. But unlike a chimeric
antibody, the variable region may be partially derived from a
human. The nonhuman, synthetic portions of a humanized antibody
often come from CDRs in murine antibodies. In any event, these
regions are crucial to allow the antibody to recognize and bind to
a specific antigen.
[0058] As noted, murine antibodies play an important role in
antibody technology. While useful for diagnostics and short-term
therapies, murine antibodies cannot be administered to people
long-term without increasing the risk of a deleterious immunogenic
response. This response, called Human Anti-Mouse Antibody (HAMA),
occurs when a human immune system recognizes the murine antibody as
foreign and attacks it. A HAMA response can cause toxic shock or
even death. Chimeric and humanized antibodies reduce the likelihood
of a HAMA response by minimizing the nonhuman portions of
administered antibodies. Furthermore, chimeric and humanized
antibodies have the additional benefit of activating secondary
human immune responses, such as antibody dependent cellular
cytotoxicity.
[0059] An "intact" antibody is one that comprises an
antigen-binding variable region as well as a light chain constant
domain (CL) and heavy chain constant domains, CH1, CH2 and CH3. The
constant domains may be native sequence constant domains (e.g.,
human native sequence constant domains) or amino acid sequence
variant thereof. The intact antibody may have one or more "effector
functions" which refer to those biological activities attributable
to the Fc region (a native sequence Fc region or amino acid
sequence variant Fc region) of an antibody. Examples of antibody
effector functions include C1q binding; complement dependent
cytotoxicity; Fc receptor binding; antibody-dependent cell-mediated
cytotoxicity (ADCC); phagocytosis; down regulation of cell surface
receptors (e.g., B cell receptor; BCR), etc.
[0060] Depending on the amino acid sequence of the constant domain
of their heavy chains, intact antibodies can be assigned to
different "classes." There are five major classes of intact
antibodies: IgA, IgD, IgE, IgG, and IgM, and several of these may
be further divided into "subclasses" (isotypes), e.g., IgG1, IgG2,
IgG3, IgG4, IgA, and IgA2. The heavy-chain constant domains that
correspond to the different classes of antibodies are called
.alpha., .delta., .epsilon., .gamma., and .mu., respectively. The
subunit structures and three-dimensional configurations of
different classes of immunoglobulins are well known.
[0061] In embodiments, the first ultrafiltering concentrates the
first antibody preparation to provide the second antibody
preparation having an antibody concentration of about 30 grams per
liter, and the second ultrafiltering concentrates the intermediate
antibody preparation (obtained from diafiltering) to provide the
third antibody preparation having an antibody concentration of, for
example, about 170 to about 200 grams per liter. The first
ultrafiltering and the second ultrafiltering can be accomplished
with the same ultra-filter membrane, and if desired, within the
same vessel or process circuit, for example, to minimize handling,
losses, leakage, and like impacts on yield, efficiency, and
economy. The first ultrafiltering and the second ultrafiltering can
be accomplished with any suitable ultrafilter apparatus or
ultrafilter membrane. Many suitable ultrafilter apparatus and
ultrafilter membranes, which are capable of tangential flow
filtration (TFF) operation to accomplish the ultrafiltrations and
diafiltration, are commercially available, such as from Millipore,
Pall Corp., Sartorius, and like vendors. In embodiments, a suitable
ultra-filter membrane can be, for example, any regenerated
cellulose composite, which composite has a relatively low protein
adsorption profile compared to other available ultra-filter
membranes, such as, polyethersulfone.
[0062] The diafiltering operation exchanges a first buffer
composition present in the first and second antibody preparations
for a second buffer desired in the third antibody preparation. In
embodiments, the first buffer can comprise, for example, a mixture
of aqueous sodium chloride and a TRIS buffer, and the second buffer
can comprise, for example, a mixture of aqueous histidine chloride
and arginine chloride. The diafiltering can accomplish a buffer
exchange at constant volume, constant concentration, or both. In
embodiments, the diafiltering accomplishes a buffer exchange at
constant volume and constant concentration. The diafiltering can
accomplish a buffer exchange, for example, of from about 5 to about
15 fold volumes (i.e. diavolumes). The diafiltering can also
accomplish a buffer exchange, for example, of about 8-fold volumes
(8 diavolumes), that is, 8 times the volume of the sample
containing the antibody preparation to be exchanged. For example, a
10 liter antibody preparation can be diafiltered with a 5 fold
(diavolumes) or 50 liter volume of exchange buffer. The exchange
volume and preferences for exchange volumes considers a balance of
factors, for example, process through-put efficiencies, product
purity, governmental and customer-patient acceptability standards,
and like standards, and can depend on, for example, the
concentration and type of buffer (e.g., the first buffer) in the
first antibody preparation, and like considerations.
[0063] The first ultrafiltering, the second ultrafiltering, and the
diafiltering are preferably accomplished with tangential flow
filtration (TFF mode) across an ultra-filter membrane, and the
ultra-filter membrane is preferably the same membrane for each
step. The yield of product in the final pool (i.e., the third
antibody preparation) can be, for example, greater than about 70
weight percent, such as from about 80 to about 100 weight percent
based on the weight of antibodies in the first antibody
preparation. The yield of the third antibody preparation can be, in
embodiments, greater than about 90 weight %, in embodiments,
greater than about 95 weight percent, and in embodiments, even
greater than about 98 weight %, based on the weight of antibodies
in the first antibody preparation.
[0064] The first ultrafiltering can have a recirculation rate of,
for example, from about 50 to 1,000 mL/min, and preferably from
about 100 to 1,000 mL/min. The recirculation rate can be scaled in
accordance with the available membrane area, for example, membrane
areas of 5, 20, 200, 1,000 square feet, and like areas permit
increasingly higher recirculation rates. Thus, a suitable scaled
recirculation rate, in embodiments, can be, for example, from about
0.5 L/min/ft.sup.2 to about 5 L/min/ft.sup.2. The ultrafiltering
and diafiltrating can be accomplished, for example, at
transmembrane pressures of from about 5 to about 50 p.s.i. The
ultrafiltering and diafiltrating can be accomplished, for example,
at transmembrane pressures of from about 10 to about 50 p.s.i. In
embodiments of the present disclosure there is provided a process
for preparing an antibody concentrate for a more dilute antibody
formulation, the antibody concentrate having a minimum bio-burden,
for example, of less than or under a detectable limit, such as,
less than about 100 CFU/mL.
[0065] Antibody compositions of the disclosure can be, for example,
concentrated monoclonal antibody preparation for administration to
humans, such as at a concentration of greater than or equal to
about 100 g/L (mg/mL), such as about 120 to about 170 g/L.
[0066] The antibody compositions of the disclosure can be, for
example, immunoglobulins, such as from the group IgA, IgD, IgE,
IgG, and IgM; sub-classes thereof; recombinants thereof; fragments
thereof; and mixtures thereof of any of the foregoing. A preferred
antibody composition of the disclosure includes recombinant
humanized anti-IgE antibodies. The antibody compositions of the
disclosure can include a buffer. A preferred buffer can be, for
example, a mixture of aqueous histidine chloride and arginine
chloride.
[0067] The preparative processes of the disclosure are preferably
accomplished in the same apparatus and without operator
intervention or with minimal operator intervention, for example, as
illustrated in FIG. 1.
[0068] The first antibody preparation can be provided or prepared
using a variety of chemical, physical, mechanical or
non-mechanical, or biochemical methods, such as, grinding,
ultrasonication, homogenization, enzymatic digestion, solvent
extraction, centrifugation, chromatography, and like methods, and
combinations thereof, see for example, the above mentioned R.
Hatti-Kaul et al., "Downstream Processing in Biotechnology," in
Basic Biotechnology, Chap. 9. The third antibody preparation can be
further processes, if desired, using for example, nanofiltration
(to remove, e.g., divalent ions), reverse osmosis (to remove, e.g.,
monovalent ions), and like liquid purification methods. The third
antibody preparation of the present disclosure can be packaged,
stored, or directly used. The third antibody preparation can be
further processed, if desired, using for example, additional
concentration steps, such as drying, lyophilization,
lyophilization-reconstitution, and like methods. The resulting
concentrated third antibody product can be reconstituted at a later
time, if desired, with a suitable liquid.
[0069] Referring to the figures, FIG. 1 illustrates an apparatus,
in embodiments of the present disclosure, for accomplishing the
preparative process including an ultrafiltration-difiltration
system (100) having an TFF ultra-filtration-difiltration (UF-DF)
unit (110), having a UF-DF membrane (115), which is in
communication with recirculation tank (120) which tank serves as a
main feed and retentate reservoir. In embodiments, tank (120) can
have a temperature control system comprising, for example, an
insulating jacket (125), a thermostatic or temperature controlled
heating element (126), such as a rheostat resistive heater element
or a circulating heated liquid system which includes a heater (not
shown), a flow regulator (127), such as a recirculating pump, and a
suitable heat transfer fluid, such as either water, glycols, or
mixtures thereof. All in-circuit components or component
contributing to in-circuit flow or processing, such as pipes,
valves, pumps, tanks, and like components, can be optionally
insulated or optionally adapted for external heating to maintain
close control over temperature specifications and to avoid
temperature excursions in the recirculating fluid loop within and
between filter chamber (110) and recirculation tank (120). In
embodiments, for example, when the system (100) is accomplishing
the first ultrafiltration or first ultrafiltering, such as in a
fed-batch mode, the system can include an optional feed tank (128)
which is in fluid communication with recirculation feed tank (120)
and can be used to, for example, make-up, replenish, or supplement
the depleted liquid phase from recirculation tank (120).
[0070] A pump (130) pumps feed liquid from tank (120) through the
UF/DF unit (110) and thereafter recirculates the resulting
retentate (the non-filtered or membrane excluded portion of the
feed liquid) to recirculating tank (120). A second tank (140) holds
and optionally pumps (not shown) a buffer into the main circuit
(110-120 loop) during the constant volume diafiltration. For
example, the addition rate and volume of the buffer introduced into
the main circuit is preferably at the same rate and volume at which
filtrate leaves the main circuit through membrane (115). Buffer
tank (140) can be optionally insulated with jacket (143) and can
include the equivalent of the abovementioned heating element and a
recirculating pump (not shown). An optional inert gas source (145),
such as nitrogen, or other compressed gas sources can be used, for
example, for product recovery, to pressurize the retentate return,
exclude oxygen, for flushing, for cleaning, for membrane integrity
testing, and like operations. A third tank (160) is used to collect
and recover the filtrate exiting the unit (110). Valves (150, 170)
can be used as appropriate to regulate the direction and optionally
the liquid flow rate in the system. All values and pumps can be
actuated manually, by coordinated computer control, or both. An
optional forth tank (190) and exit stream can provide an ancillary
waste-flush, product recovery, or monitoring system, for example,
when equipped with an optional monitoring device (180), such as an
optical density meter, optional filter(s) (185) such as a guard
filter, product filter, and like optional subsystems. In
embodiments, the main fluid circuit (110-120 loop) can optionally
be equipped with an in-line monitoring system.
[0071] The concentrated antibody preparations prepared by processes
of the present disclosure can be used for human therapeutic
administration, including immunoglobulin products, for either
intramuscular (IMIG) or intravenous (IVIG) administration. The
concentrated antibody preparations of the disclosure can include a
stabilizer, for example, a buffered amino acid salt solution,
simple sugars, or like stabilizers, suitable ions chelators, such
as EDTA or citrate ion, and combinations thereof, see for example,
Wang, Y.-C. J. et al, "Parenteral formulations of proteins and
peptides: stability and stabilizers," J. Parenteral Sci. Technol.,
42, Suppl. S3-S26 (1988). Derwent Abstract of JP01268646A
(AN89-359879) reports that the application describes an injection
preparation of an IgG.sub.3 monoclonal antibody having a
concentration of 0.1 micrograms/mL to 100 mg/mL. Subject matter
disclosed in these publications is believed to be outside the scope
of the present disclosure.
[0072] Preparations according to the disclosure can be
substantially free from aggregates. Acceptable levels of aggregated
contaminants would be less than, for example, about 5 weight %, and
ideally less than 2 weight %. Levels as low as 0.2 weight % can be
achieved, although aggregated contaminants of about 1 weight % is
more typical. The preparation in embodiments, can also preferably
be free from excipients traditionally used to stabilize polyclonal
formulations, for example glycine and/or maltose.
[0073] The present disclosure can provide a monoclonal antibody
preparation for administration to a human characterized in that the
antibody in the preparation is a recombinant antibody and can be at
a concentration of 100 mg/mL or greater, preferably greater than
150 mg/mL. The preparation is preferably substantially free from of
any protein aggregation.
[0074] The pH of pharmaceutical formulations of the disclosure will
depend upon the particular route of administration. However, in
order to maximize the solubility of the antibody in the
concentrated solution, the pH of the solution should be different
from the pH of the isoelectric point (pI) of the antibody.
[0075] In embodiments of the disclosure, the monoclonal preparation
can be envisaged for use in human therapy. Various human disorders
can be treated such as cancer or infectious diseases, for example,
those mentioned above, and immune dysfunction such as
T-cell-mediated disorders including severe vasculitis, rheumatoid
arthritis, systemic lupus, also autoimmune disorders such as
multiple sclerosis, graft versus host disease, psoriasis, juvenile
onset diabetes, Sjogrens' disease, thyroid disease, myasthenia
gravis, transplant rejection, inflammatory bowel disease, asthma,
IgE mediated disorders, and like disorders or conditions, or
combinations thereof.
[0076] The disclosure therefore provides in embodiments the use of
a concentrated monoclonal antibody preparation as described herein
in the manufacture of medicament for the treatment of any of the
aforementioned disorders, and like disorders. Also provided is a
method of treating a human being, having any such disorder,
comprising administering to the individual a therapeutically
effective amount of a preparation according to the disclosure. The
dosages of such antibody preparations will vary with the conditions
being treated and the recipient of the treatment, but can be, for
example, in the range of about 50 to about 2,000 mg for an adult
patient preferably about 100 to about 1,000 mg administered daily
or weekly for a period between 1 and 30 days, and repeated as
necessary. The doses may be administered as single or multiple
doses.
Process Description. The formulation step typically exchanges the
purified bulk drug substance, for example, resulting from
ion-exchange chromatography, into the final excipient composition
and concentration. There was typically no purification achieved at
this step except for small molecule removal. The emphasis was on
high yield, buffer exchange, and formulation step robustness.
During formulation via TFF (tangential flow filtration), the
protein-containing feed solution was pumped through the membrane
system and back to the recycle (recirculation) vessel. The TFF
membrane retained the protein (as part of the retentate) while the
filtrate (or permeate) was driven through the membrane by pressure.
The pressure is called the transmembrane pressure (TMP) and is
typically controlled using a retentate pressure control valve. The
process was usually achieved by a sequence of a first
ultrafiltering (concentration), diafiltering (constant volume
buffer exchange), and a second ultrafiltering (further
concentration). The number of diavolumes (volumetric equivalents)
necessary to remove process buffer components can be readily
calculated or determined experimentally. UF/DF Process Generally
for anti-IgE. The pH of an anion-exchange pool from chromatography
was adjusted to a pH of about 6 using 0.5 M aqueous phosphoric
acid. The pH adjusted anion-exchange pool was formulated by
ultrafiltration/diafiltration (UF/DF) process of the present
disclosure using a membrane having a nominal molecular cut off of
10,000-30,000 Daltons. Prior to processing, the UF membrane was
equilibrated with diafiltration buffer (0.02 M histidine, 0.2 M
arginine-HCl, pH 6).
[0077] The product from an anionic exchange (anion-exchange pool)
was then loaded on the system and was concentrated to an
intermediate concentration by the first ultrafiltering. The pool
was then diafiltered (8.times. or diavolumes) into its formulation
(0.02 M histidine, 0.2 M arginine-HCl, pH 6). The pool was then
concentrated by a second ultrafiltering to a final bulk
concentration of >170 g/L and recovered through a 0.22
micrometer sterile filter. The entire UF/DF process was performed
at a temperature set point of about 45 degrees C. This temperature
control was achieved using temperature control of the incoming
anion-exchange pool, the diafiltration buffer, and the use of
ajacketed recirculation vessel for the UF/DF process as illustrated
herein.
[0078] After UF/DF, the recovered pool was diluted (i.e.,
conditioned) to a bulk concentration of about 150 g/L in 0.02 M
histidine, 0.2 M arginine-HCl, 0.04% polysorbate-20, pH 6 (final
formulation). During the conditioning steps the temperature of the
bulk was allowed to return to ambient temperature. After
conditioning, the formulated bulk was again recovered through a
0.22 micrometers sterile filter.
[0079] The UF/DF system can be regenerated with 0.1 N sodium
hydroxide and sanitized with 1.4% Minncare.RTM.. When not in use
the system can be stored in 0.1N aqueous sodium hydroxide. The
UF/DF membranes can be stored, for example, in a 0.1%
Roccal.RTM./20% glycerol-water solution between campaigns.
General Ultrafiltration/Diafiltration Process Procedures
[0080] Operating Parameters Feed flow rate at 0.5 L/min/ft.sup.2. A
constant retentate pressure (e.g., 10 psig) control was used for
cleaning and pre-use equilibration, whereas C.sub.wall, constant
retentate pressure or constant TMP was used for processing.
[0081] Pre-Use Equilibration: The following preparations were
accomplished on cleaned Pellicon-2 cassette membranes prior to use
to assure the membranes were properly equilibrated.
TABLE-US-00001 Volume (L/ft.sup.2) Solution (room temp) Mode -- --
SPFO 1.0 WFI SPFO 1.0 DF buffer SPFO 0.5 DF Buffer TRFO, 10 minutes
-- -- SPFO
[0082] Process Use: The following was performed on the resulting
initial anion-exchange pool (Q-pool) obtained from a preceding
separation step, for example, a Q-Sepharose chromatography
step:
[0083] a first ultrafiltering or first ultrafiltration (UF1) to a
concentration from about 5 g/L to a concentration for difiltration
(C.sub.DF);
[0084] diafiltering or diafiltration (DF1) with four (4)
difiltration volumes (DV) with the DF buffer;
[0085] continued diafiltering (DF2) with four (4) difiltration
volumes (DV) of DF buffer;
[0086] a second ultrafiltering or second ultrafiltration (UF2) to a
final concentration (C.sub.Final); and
[0087] optional product recovery.
[0088] The foregoing steps were typically accomplished at low dP
Recycle (mix), for example, 15 min.
[0089] Post-Use Cleanout: The following tabulated sequence and
conditions were used for cleanout on the Pellicon-2 cassette
membranes immediately following use.
TABLE-US-00002 Volume Solution (L/ft.sup.2) (room temp) Mode 1.0
0.1N NaOH SPFO 0.5 0.1N NaOH TRFO, 30 minutes -- -- SPFO 1.0 WFI
SPFO 0.5 300 ppm TRFO, 30 Minncare .RTM. minutes -- -- SPFO 1.0 WFI
SPFO -- -- Integrity Test @ 30 psig 0.5 0.1N NaOH TRFO, 15 minutes
storage
DEFINITIONS FOR MODES OF OPERATION IN TFF
[0090] Single Pass with Filtrate Open (SPFO). The retentate and
filtrate are directed to drain. Filtrate valve open.
[0091] Total Recycle with Filtrate Open (TRFO). The retentate and
filtrate are directed to recycle vessel. Filtrate valve open.
[0092] Fed-Batch Ultrafiltration (FB-UF). The retentate is directed
to the recycle tank, the filtrate directed to drain, and the
incoming pool transferred into the recycle tank.
[0093] Batch Ultrafiltration (B-UF). The retentate is directed to
the recycle tank and the filtrate is directed to drain.
[0094] Diafiltration (DF). The retentate is directed to the recycle
tank, the filtrate is directed to drain, and the diafiltration
buffer is transferred into recycle tank.
[0095] dP refers to differential pressure.
[0096] Product Transfer. The ultrafilter membrane unit and recycle
tank are open to the pool tank. The nitrogen overlay pressure is
controlled. The pool is transferred first using the recycle pump
and then using a manual peristaltic pump.
[0097] Feed Transfer. The incoming pool is pumped into the recycle
tank.
[0098] Total Recycle with Filtrate Closed (TRFC). The retentate is
directed to a recycle vessel. Filtrate valve closed.
[0099] "Q-pool" refers to the protein pool resulting from, for
example, a preceding Q-Sepharose chromatography step which has been
conditioned with buffer, also referred to as the "conditioned
pool."
[0100] WFI refers to water-for-injection.
EXAMPLES
[0101] The following examples serve to more fully describe the
manner of using the above-described disclosure, as well as to set
forth the best modes contemplated for carrying out various aspects
of the disclosure. It is understood that these examples in no way
serve to limit the true scope of this disclosure, but rather are
presented for illustrative purposes.
Example 1
[0102] High Concentration Formulation of rhuMAb E25 A pilot scale
UF system was used to concentrate/formulate rhuMAb E25 (a
recombinant human monoclonal antibody that targets IgE, U.S. Pat.
No. 6,172,213). A Millipore Pelicon Ultrafiltration/Diafiltration
system was assembled with a 5.7-sqft, 10,000-dalton regenerated
cellulose composite membrane. The system consisted of a membrane
holder, a Waukeskaw Model 6 rotary lobe feed pump, 1/2'' 316 L
stainless steel recirculation piping, and a recirculation vessel.
Pressure indicators/transmitters (Anderson) were located at the
inlet (FEED), outlet (RETENTATE) and permeate (FILTRATE) of the
membrane holder. Flow meters (Yokogawa ADMAG) were located at the
inlet (FEED) and permeate (FILTRATE) of the membrane holder. A
back-pressure regulating valve (Mikroseal) was located at the
outlet of the membrane holder to control the retentate pressure and
effect the transmembrane pressure (TMP). A 40-liter 316 L stainless
steel jacketed tank was used for the recirculation vessel. This
tank was fitted with a level indicator, top-mounted agitator
(Lightnin), vortex breaker and bottom valve (NovAseptic).
Temperature control was achieved through the use of temperature
modulated glycol fed to the jacket of the tank.
[0103] During this run, the feed flow rate was set to a constant
rate of 2.85 L/min (0.5 L/min/ft.sup.2). During all pre-use and
post-use operations the retentate pressure control was set to a
constant of 10 psig. During the ultrafiltering and diafiltering
operations the system used a C.sub.wall control scheme to control
the flux through the membrane, see for example R. van Reis, et al.,
Constant C.sub.wall Ultrafiltration Process Control, J. of Membrane
Science, 130 (1997), 123-140.
[0104] Prior to the process, the system storage solution (0.1N
NaOH) was flushed in a single pass to drain mode, first with 2
L/ft.sup.2 purified water (PW) and then 1 L/ft.sup.2 diafiltration
buffer (50 mM Histidine/pH 6.0). After the flushes, the system was
equilibrated by recirculating 0.5 L/ft.sup.2 diafiltration buffer
for 10 min. The pH of the recirculated solution was checked to
confirm the equilibration. The level in the tank was then reduced
to a minimum measurable value to minimize dilution of the incoming
protein pool. The protein pool resulting from a preceding
Q-Sepharose chromatography step was measured to be 3.2 g E25/L and
had a volume of 43.1 L. The protein was in a solution of 25 mM TRIS
buffer and about 200 mM NaCl and pH adjusted to 6.2. To begin the
run the protein pool was transferred to the recirculation vessel.
In the vessel the pool was agitated via the top mounted impeller
and the temperature was maintained at ambient (20-25.degree.
C.).
[0105] During the process the pool was concentrated in UF1 mode to
50 g E25/L (about 2.8 L). At the beginning of diafiltration the
temperature set point of the recirculation vessel was increased to
40.degree. C. The increase in temperature and control was affected
by flowing warm glycol through the outer jacket of the tank. The
pool was then diafiltered with 8 diavolumes of diafiltration
buffer. The diafiltration was performed at a constant volume, which
was achieved by matching the flow rate of buffered solution being
transferred into the recirculation tank to the flow rate of
filtrate being removed from the system. At the end of the
diafiltration, the pool was further concentrated in UF2 mode. This
phase was also performed using an elevated temperature set point of
40.degree. C. The target for this final concentration was 110 g/L.
This was achieved without the need to reduce the feed flow rate.
Next, a low pressure drop mixing was performed where the feed pump
was controlled to maintain a 5-10 psig pressure drop across the
feed channel. A sample was pulled from the recirculation tank and a
final bulk concentration of approximately 120 g/L was measured.
Table 1 summarizes the throughput and flux results of UF1, DF
(DF1+DF2), and UF2.
TABLE-US-00003 TABLE 1 Normalized Throughput Normalized Flux
Process Phase (g/ft.sup.2/hr) (LMH/psig) UF1 13.8 4.97 DF 13.8 2.92
UF2 181.4 1.64
[0106] FIG. 2 shows the observed or measured process values over
time for the feed flow rate (210), tank temperature (220), fed dP
(230), TMP (240), and filtrate flow rate (250) parameters during
the various phases or mode of the process including UF1 (10), DF
(20), UF2 (30).
[0107] FIG. 3 shows the observed or measured process values over
time for the E25 concentration (310), flux (320), and TMP
(240).
[0108] FIG. 4 shows the observed or measured process values over
time for pressure drop versus protein concentration observed for
UF1 (410) and UF2 (420) at 37.degree. C.
[0109] The protein pool was recovered by a series of steps. First
the pool in the recirculation tank was pumped from the tank through
a Millipac 200, 0.22 microns sterilizing grade filter using the
rotary lobe feed pump. Next the protein solution was displaced from
the piping and membrane unit with a 5 psig nitrogen gas blow down
applied to the highest point on the retentate line. The final phase
was a blow down of the tank and feed line, also using the 5 psig
nitrogen gas.
[0110] The product recovery was believed to be improved compared to
Example 1 when conducted at ambient temperature because the
elevated temperature used in one or more of the ultrafiltering,
diafiltering, or recovery steps reduced viscous effects. For
example, when the temperature control was turned off during product
recovery, the system slowly cooled during this operation causing
difficulties for recovery from the membrane unit. Alternatively,
the recovery can be performed first from the membrane holder and
then from the recirculation vessel.
[0111] To determine the mass of loss during recovery, 1.74 L of DF
buffer was added to the system and recirculated for about 5 minutes
and recovered using the same sequence as described above. This
volume was then analyzed for protein concentration with the other
pools. Table 2 summarizes the results.
TABLE-US-00004 TABLE 2 Concentration Yield or Volume (L) (g/L) Mass
(g) {Loss} (%) Q-Pool 43.1 3.2 137.9 100 Recovered 0.99 120 118.8
86.1 Pool Buffer 1.74 9.8 17.1 12.4 Flush Filtrate 65.3 0.04 2.6
1.9
[0112] Post processing, the membrane was regenerated using 0.1N
NaOH, 1 L/ft.sup.2 single pass flush followed by 0.5 L/ft.sup.2
total recirculation for 30 min. This was followed by 1 L/ft.sup.2
PW (pure water) flush. This was followed by a total recirculation
of 300 ppm Mirncare.RTM. solution for 30 min. The system was again
flushed with 1 L/ft.sup.2 PW and finally recirculated for 15 min
with 0.1N NaOH and stored. The recovered pool was diluted to 80 g
E25/L and conditioned into the final formulation of 50 mM
histidine/150 mM trehalose/0.02% polysorbate 20/pH 6.0. Product
quality was assessed by size exclusion chromatography (SEC) for
both the incoming Q-Pool and final recovered bulk. This data is
summarized in Table 3.
TABLE-US-00005 TABLE 3 SEC Results (% Pool monomer) Q-Pool 99.8
Final Bulk 99.8
Comparative Example 2
[0113] High Concentration Formulation of rhuMAb E25 at Ambient
Temperature Example 1 was accomplished with the following
exceptions. Prior to the process, the system storage solution (0.1N
NaOH) was flushed in a single pass to drain mode first with 2
L/ft.sup.2 purified water (PW) and then 1 L/ft.sup.2 diafiltration
buffer (20 mM histidine/pH 6.0). After the flushes the system was
equilibrated by recirculating 0.5 L/ft.sup.2 diafiltration buffer
for 10 min. The pH of the recirculated solution was checked to
confirm the equilibration. The level in the tank was then reduced
to a minimum measurable value to minimize dilution of the incoming
protein pool.
[0114] The protein pool resulting from the preceding Q-Sepharose
chromatography step was measured to be 3.3 g E25/L and had a volume
of 33.3 L. The protein was in a solution of 25 mM TRIS buffer and
about 200 mM NaCl and pH adjusted to 6.2. To begin the run the
protein pool was transferred to the recirculation vessel. In the
vessel the pool was agitated via the top mounted impeller and the
temperature was maintained at ambient (20-25.degree. C.). During
the process the pool was concentrated down in UF1 mode to 50 g
E25/L (about 2.2 L). The pool was then diafiltered with 8
diavolumes of diafiltration buffer. The diafiltration was performed
at a constant volume, which volume was achieved by matching the
flow rate of buffered being transferred into the recirculation tank
to the flow rate of filtrate being removed from the system. The
diafiltration was also performed at ambient temperature. At the end
of the diafiltration, the pool was further concentrated in UF2
mode. The target for this final concentration was 110 g/L. However,
due to a high pressure drop across the feed channel, this
concentration was not achieved. In an attempt to achieve this
concentration the feed flow rate was reduced to 1.4 L/min at a bulk
concentration of about 80 g E25/L because the pressure drop across
the feed channel had reached 50 psig. UF2 was continued until a
high pressure drop of 50 psig again was reached and the process was
stopped. Next, a low pressure drop mixing was attempted where the
feed pump was used to maintain a 5 psig pressure drop across the
feed channel. Again, the viscous nature of the protein solution
made it difficult to achieve since the rotary lobe pump reached
excess pressures. A sample was pulled from the recirculation tank
and a final bulk concentration of approximately 104 g/L was
measured. Table 4 summarizes throughput and flux measured during
the UF1, DF (DF1+DF2), and UF2 phases.
TABLE-US-00006 TABLE 4 Normalized Process Throughout Normalized
Flux Phase (g/ft.sup.2/hr) (LMH/psig) UF1 14.5 5.31 DF 9.5 1.47 UF2
144.6 0.78
[0115] FIG. 5 shows the observed or measured process values over
time for the feed flow rate (210), tank temperature (220), fed dP
(230), TMP (240), and filtrate flow rate (250) parameters during
the various phases or mode of the process including UF1 (10), DF
(20), UF2 (30).
[0116] FIG. 6 shows the observed or measured process values over
time for the E25 concentration (310), flux (320), and TMP
(240).
[0117] FIG. 7 shows the observed or measured process values over
time for pressure drop versus protein concentration observed for
UF1 (410) and UF2 (420) at 24.degree. C.
[0118] The protein pool was recovered in steps. First, the pool in
the recirculation tank was pumped from the tank through a Millipac
200, 0.22 microns sterilizing grade filter using the rotary lobe
feed pump. Next the protein solution was displaced from the piping
and membrane unit with a 5 psig nitrogen gas blow down applied to
the highest point on the retentate line. The product recovery from
this was very poor due to the viscous nature of the solution. The
final phase was a blow down of the tank and feed line, also using
the 5 psig nitrogen gas.
[0119] To determine the mass of loss during recovery, 1.85 L of DF
buffer was added to the system and recirculated for about 5 minutes
and recovered using the sequence of Example 1. This volume was then
analyzed for protein concentration with the other pools. Table 5
summarizes the results.
TABLE-US-00007 TABLE 5 Concentration Yield or Volume (L) (g/L) Mass
(g) {Loss} (%) Q-Pool 33.3 3.3 109.9 100 Recovered 0.77 104.4 80.4
73.1 Pool Buffer 1.85 14.7 27.2 24.7 Flush Filtrate 52.2 0.03 1.6
1.5
[0120] Post process, the membrane was regenerated using 0.1N NaOH,
1 L/ft.sup.2 single pass flush followed by 0.5 L/ft.sup.2 total
recirculation for 30 min. This was followed by 1 L/ft2 PW flush.
This was followed by a total recirculation of 300 ppm Minncare.RTM.
solution for 30 min. The system was again flushed with 1 L/ft.sup.2
PW and finally recirculated for 15 min with 0.1N NaOH and stored.
The recovered pool was diluted to 80 g E25/L and conditioned into
the final formulation of 20 mM histidine/250 mM sucrose/0.02%
polysorbate 20/pH 6.0. Product quality was assessed by size
exclusion chromatography (SEC) for both the incoming Q-Pool and
final recovered bulk. This data is summarized in Table 6.
TABLE-US-00008 TABLE 6 SEC Results (% Pool monomer) Q-Pool 99.8
Final 99.8 Bulk
Example 3
[0121] High Concentration Formulation of rhuMAb E26 with Initial
Fed-Batch Mode Example 1 was repeated with the following
exceptions. The concentrate/formula was rhuMAb E26 (a recombinant
human monoclonal antibody that targets IgE). The products from this
example were used in toxicology assessment. The Millipore Pelicon
Ultrafiltration/Diafiltration system was assembled with a 11.4-sqft
30,000-Dalton regenerated cellulose composite membrane. The feed
flow rate was set to a constant rate of 5.0 L/min (0.44
L/min/ft.sup.2). During the ultrafiltration and diafiltration
operations the retentate pressure was maintained between about 6-8
psig. The protein pool resulting from the preceding Q-Sepharose
chromatography step was measured to be 6.7 g E26/L and had a volume
of 59.3 L.
[0122] Because the incoming pool was larger then the recirculation
vessel, the UF1 process began in fed-batch mode. In this mode,
Q-Pool was added to the recirculation vessel at approximately the
same rate as filtrate passes through the TFF membrane to drain.
After the remaining Q-Pool had transferred to the recirculation
vessel, the UF1 process continued in batch mode. During the UF1 the
pool was concentrated to 50 g E26/L (about 7.9 L). At the beginring
of diafiltration the temperature set point of the recirculation
vessel was increased to 40.degree. C. The increase in temperature
and control was affected by flowing warm glycol through the outer
jacket of the tank. The pool was then diafiltered with 8 diavolumes
of diafiltration buffer. The diafiltration was performed at a
constant volume which was achieved by matching the flow rate of
buffered being transferred into the recirculation tank to the flow
rate of filtrate being removed from the system. At the end of the
diafiltration, the pool was further concentrated in UF2 mode to a
final concentration of 109 g E26/L (3.6 L). This phase was also
performed using an elevated temperature set point of 40.degree. C.
Next a low pressure drop mixing was performed where the feed pump
was controlled to maintain a 5-10 psig pressure drop across the
feed channel. Table 7 summarizes the throughput and flux results of
UF1, DF (DF1+DF2), and UF2.
TABLE-US-00009 TABLE 7 Normalized Process throughput Normalized
Flux Phase (g/ft.sup.2/hr) (LMH/psig) UF1 26.1 3.71 DF 19.2 2.34
UF2 174.2 1.80
[0123] FIG. 8 shows the observed or measured process values over
time for the feed flow rate (210), tank temperature (220), fed dP
(230), TMP (240), and filtrate flow rate (250).
[0124] FIG. 9 shows the observed or measured process values over
time for the E26 concentration (910), flux (920), and TMP
(940).
[0125] FIG. 10 shows the observed or measured process values over
time for pressure drop versus protein concentration observed for
UF1 (1010) and UF2 (1020).
[0126] Just prior to product recovery, a 10 mL sample was analyzed
for detection and a titer of bioburden. A typical reject limit is
1,000 Colony Forming Units (CFU) per mL. The results of this test
were 1.8 CFU/mL, a suitable value at this step and well below the
reject limit. To determine the mass of loss during recovery, 908.1
mL of DF buffer was added to the system and recirculated for about
5 minutes and recovered using the same sequence described above.
This volume was then analyzed for protein concentration with the
other pools. Table 8 summarizes the results.
TABLE-US-00010 TABLE 8 Concentration Yield or Volume (L) (g/L) Mass
(g) {Loss} (%) Q-Pool 59.3 6.7 397.3 100 Recovered 3.41 109.1 372.0
93.6 Pool Buffer 0.908 20.4 18.5 4.7 Flush Filtrate 120 n/d n/d
n/d
[0127] The recovered pool was diluted to 80 g E26/L and conditioned
into the final formulation of 50 mM histidine/150 mm
trehalose/0.02% polysorbate 20/pH 6.0. Product quality was assessed
by size exclusion chromatography (SEC) for the incoming Q-Pool, the
retentate pool after UF1, the retentate pool after DF, and final
recovered bulk. This data is summarized in Table 9.
TABLE-US-00011 TABLE 9 SEC Results Pool (% monomer) Q-Pool 99.8 End
of 99.8 UF1 End of 99.8 DF Final 99.8 Bulk
Example 4
[0128] High Concentration Formulation of rhuMAb E26 for Toxicology
Evaluation-Comparison of 10 kD and 30 kD Example 3 was repeated
with the following exceptions. Two pilot scale UF systems were used
to concentrate/formulate rhuMAb E26. Two Millipore Pelicon
Ultrafiltration/Diafiltration systems were assembled with a
11.4-sqft, regenerated cellulose composite membrane, one with
10,000-Dalton pore size and the other a 30,000-Dalton pore size.
The retentate pressures were maintained at about 6-9 psig.
10 kD Process
[0129] The protein pool resulting from the preceding Q-Sepharose
chromatography step was measured to be 5.85 g E26/L and had a
volume of 62.4 L. During the UF1, the pool was concentrated to 50 g
E26/L (about 7.3 L). At the end of the diafiltration, the pool was
further concentrated in UF2 mode to a final concentration of 107.5
g E26/L (3.4 L). Table 10 summarizes the throughput and flux
results of UF1, DF, and UF2.
TABLE-US-00012 TABLE 10 Process Normalized Normalized Flux Phase
throughput (g/ft.sup.2/hr) (LMH/psig) UF1 21.8 3.6 DF 15.9 2.6 UF2
137.4 1.93
[0130] To determine the mass of loss during recovery, 987 mL of DF
buffer was added to the system and recirculated for about 5 minutes
and recovered using the same sequence described above. This volume
was then analyzed for protein concentration with the other pools.
Table 11 summarizes the results.
TABLE-US-00013 TABLE 11 Concentration Yield or Volume (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 62.4 5.85 365.4 100 Recovered 3.38 107.5
361.7 98.9 Pool Buffer 0.987 19.9 19.6 5.4 Flush Filtrate 125 n/d
n/d n/d
[0131] FIG. 11 shows the observed or measured process values over
time for the feed flow rate (210), tank temperature (220), fed dP
(230), TMP (240), and filtrate flow rate (250) over the various
phases or mode of the process including UF1 (10), DF (20), UF2
(30), and low dP (40), for the 10 kD process.
[0132] FIG. 12 shows the observed or measured process values over
time for the E26 concentration (1210), flux (1220), and TMP (1240)
over the various phases or mode of the process including UF1 (10),
DF (20), UF2 (30), and low dP (40), for the 10 kD process.
[0133] FIG. 13 shows the observed or measured process values over
time for pressure drop versus protein concentration observed for
UF1 (1310) and UF2 (1320) for the 10 kD process.
30 kD Process
[0134] The protein pool resulting from the preceding Q-Sepharose
chromatography step was measured to be 5.85 g E26/L and had a
volume of 64.5 L. During the UF1 the initial pool was concentrated
to 50 g E26/L (about 7.5 L). At the end of the diafiltration, the
pool was further concentrated in UF2 mode to a final concentration
of 117.5 g E26/L (3.2 L). Table 12 summarizes the throughput and
flux results of UF1, DF, and UF2.
TABLE-US-00014 TABLE 12 Normalized Process throughput Normalized
Flux Phase (g/ft.sup.2/hr) (LMH/psig) UF1 25.5 4.01 DF 17.6 2.39
UF2 180.5 1.57
[0135] To determine the mass of loss during recovery, 918 mL of DF
buffer was added to the system and recirculated for about 5 minutes
and recovered using the same sequence described above. The
recovered pool was diluted to 80 g E26/L and conditioned into the
final formulation of 50 mM histidine/150 mM trehalose/0.02%
polysorbate 20/pH 6.0. Table 13 summarizes the results.
TABLE-US-00015 TABLE 13 Concentration Yield or Volume (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 64.5 5.85 377.3 100 Recovered 3.20 117.5
376.0 99.6 Pool Buffer 0.918 22.7 20.8 5.5 Flush Filtrate 125 n/d
n/d n/d
[0136] FIG. 14 shows the observed or measured process values over
time for the feed flow rate (210), tank temperature (220), fed dP
(230), TMP (240), and filtrate flow rate (250) over the various
phases or mode of the process including UF1 (10), DF (20), UF2
(30), and low dP (40), for the 30 kD process.
[0137] FIG. 15 shows the observed or measured process values over
time for the E26 concentration (1510), flux (1520), and TMP (1540)
over the various phases or mode of the process including UF1 (10),
DF (20), UF2 (30), and low dP (40), for the 30 kD process.
[0138] FIG. 16 shows the observed or measured process values over
time for pressure drop versus protein concentration observed for
UF1 (1610) and UF2 (1620) for the 30 kD process.
Example 5
[0139] Liquid rhuMAb E25 Scale Up Example 1 was repeated with the
following exceptions. A production scale UF system was used to
concentrate/formulate a liquid rhuMAb E25 (a recombinant human
monoclonal antibody that targets IgE). The product can be used in
therapeutic application and human bio-equivalency trials. The
Millipore Pelicon Ultrafiltration/Diafiltration systems were
assembled with a 226-sqft regenerated cellulose composite membrane,
with a pore size of 30,000-Dalton. Each system consisted of a
membrane holder, a Viking S3S rotary lobe feed pump, 11/2'' 316 L
stainless steel recirculation piping, and a 250-L recirculation
vessel.
[0140] One 250-liter 316 L stainless steel jacketed tank was used
for the recirculation vessel. Temperature control to this tank was
achieved with a temperature modulated glycol fed to the tank's
jacket. The temperature of the glycol fed to the tank jacket was
raised or lowered using either steam-fed heat exchanger or cold
glycol supply respectively.
[0141] For this run, the feed flow rate was set to a constant rate
of 114 L/min (0.5 L/min/ft.sup.2). The diafiltration buffer (20 mM
histidine/200 mM arginine chloride/pH 6.0) was prepared in a
separate tank. The temperature of this buffer was set to 45.degree.
C. prior to the process. This enabled accurate temperature control
throughout the process.
[0142] Prior to processing, the system storage solution (0.1N NaOH)
was flushed in a single pass to drain mode first with 1 L/ft.sup.2
water for injection (WFI) and then 1 L/ft.sup.2 diafiltration
buffer. After the flushes, the system was equilibrated by
recirculating 0.5 L/ft.sup.2 diafiltration buffer for 10 min. The
pH of the recirculated solution was checked to confirm the
equilibration.
[0143] The protein pool resulting from the preceding Q-Sepharose
chromatography step was measured to be 5.2562 g E25/L and had a
volume of 1,141 L. The protein was in a solution of 25 mM TRIS
buffer and about 200 mM NaCl and the pH was adjusted to 6.2. Just
prior to the run, the temperature set point of this pool was set to
45.degree. C. To begin the run the protein pool was transferred to
the recirculation vessel, through a 0.22 microns sterilizing grade
filter to a level of about 200 L in the tank. In the vessel the
pool was agitated via a top mounted impeller and the temperature
was maintained at about (40-50.degree. C.). Because the incoming
pool was larger then the recirculation vessel, the UF1 process
began in fed-batch mode. In this mode, Q-Pool was added to the
recirculation vessel at approximately the same rate as filtrate
passes through the TFF membrane to drain. After the remaining
Q-Pool was transferred to the recirculation vessel, the UF2 process
was continued in batch mode. During the UF1 mode the pool was
concentrated to about 30 g E25/L (about 200 L). The pool was then
diafiltered with about 8 diavolumes of diafiltration buffer. During
diafiltration the temperature was maintained between 40.degree. and
50.degree. C. The diafiltration was performed at a constant volume,
which was achieved by matching the flow rate of buffer being
transferred into the recirculation tank to the flow rate of
filtrate being removed from the system. At the end of the
diafiltration, the pool was further concentrated in UF2 mode to a
final concentration set point of >170 g E25/L (35 L). This UF2
mode phase was also performed at an elevated temperature set point
of 45.degree. C.+/-5.degree. C. Next, a low pressure drop mixing
was performed where the feed pump was controlled to maintain a 5-10
psig pressure drop across the feed channel. A sample was pulled and
a spec scan was performed to confirm the concentration prior to
recovery. The concentration of this sample was 219 g E25/L. Table
14 summarizes throughput and flux measured during the UF1, DF
(DF1+DF2), and UF2 phases.
TABLE-US-00016 TABLE 14 Normalized Process throughput Normalized
Flux Phase (g/ft.sup.2/hr) (LMH/psig) UF1 43.8 3.34 DF 25.9 2.46
UF2 78.9 0.66
[0144] Just prior to product recovery, a 30 mL sample was pulled
and submitted for detection and titer of bioburden. The result was
<0.13 CFU/mL. The protein pool was recovered by a series of
steps. First, the product was displaced from the membrane in a
single pass mode using 5 L of DF buffer added to the retentate
line. The product was filtered into a recovery tank through a 7.4
ft.sup.2, 0.22 microns sterilizing-grade guard filter followed by a
2 ft.sup.2, 0.22 microns sterilizing-grade final filter. The pool
in the recirculation tank was then pumped from the tank using the
rotary lobe feed pump. Next the residual protein solution was
displaced from tank and feed line with a 5 psig nitrogen gas blow
down. The final phase was a blow down of the membrane unit, which
now contained mostly DF buffer from the initial product
displacement. This phase also used the 5 psig nitrogen gas applied
to the highest point on the retentate line. The recovered pool was
diluted first to about 153 g E25/L using DF buffer. Finally, the
pool was conditioned into the final formulation of 20 mM
histidine/200 mM arginine-HCl/0.04% polysorbate 20/pH 6.0. The
volumes of the recovered pool, diluted pool, and conditioned pool
(Q-pool) were then each analyzed for protein concentration. Table
15 summarizes the results.
TABLE-US-00017 TABLE 15 Volume Concentration Yield or (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 1,141 5.2562 5,997.3 100 Recovered 35.0
170.0 5,950.0 99.2 Pool Diluted 39.0 147.0 5,726 95.5 Pool
[0145] FIG. 17 shows feed flow rate (210), tank temperature (220),
fed dP (230), TMP (240), and filtrate flow rate (250) parameters
during the various phases or mode of the process including UF1
(10), DF1 (20), DF2 (25), UF2 (30), and low dP (50).
Example 6
[0146] Liquid rhuMAb E25 Preparation Example 5 was repeated with
following exceptions. A production scale UF system was used to
concentrate/formulate liquid rhuMAb E25 (E25, a recombinant human
monoclonal antibody that targets IgE). The Millipore Pelicon
Ultrafiltration/Diafiltration systems were assembled with a
226-sqft regenerated cellulose composite membrane, with a pore size
of 30,000-dalton. Each system consisted of a membrane holder, a
Viking S3S rotary lobe feed pump, 11/2'' 316 L stainless steel
recirculation piping, and a 250-L recirculation vessel. One
250-liter 316 L stainless steel jacketed tank was used for the
recirculation vessel. The feed flow rate was set to a constant rate
of 114 L/min (0.5 L/min/ft.sup.2). During all pre-use and post-use
operations the retentate pressure control was set to a constant of
10 psig. During the ultrafiltration and diafiltration operations
the system used the C.sub.wall control scheme to control the flux
through the membrane. The diafiltration buffer (20 mM Histidine/200
mM arginine chloride/pH 6.0) was prepared in a separate tank. The
temperature of this buffer was set to 45.degree. C. prior to the
process. This enabled accurate temperature control through the
entire process. The protein pool resulting from the preceding
Q-Sepharose chromatography step was measured to be 5.5438 g E25/L
and had a volume of 1,082 L. The protein was in a solution of 25 mM
TRIS buffer and about 200 mM NaCl and pH adjusted to 6.2. Just
prior to the run, the temperature set-point of this pool was set to
45.degree. C. To begin the run the protein pool was transferred to
the recirculation vessel, through a 0.22 microns sterilizing grade
filter to a level of about 200 L in the tank. In the vessel the
pool was agitated via the top mounted impeller and the temperature
was maintained at ambient (40-50.degree. C.). Because the incoming
pool was larger then the recirculation vessel, the UF1 process
began in fed-batch mode. In this mode, Q-Pool was added to the
recirculation vessel at approximately the same rate as filtrate
passes thought the TFF membrane to drain. After the remaining
Q-Pool had transferred to the recirculation vessel, the UF1 process
continued in batch mode. During the UF1 the pool was concentrated
to about 30 g E25/L (about 200 L). The pool was then diafiltered
with 8 diavolumes of diafiltration buffer. During diafiltration the
temperature was maintained between 40 and 50.degree. C. The
diafiltration was performed at a constant volume, which was
achieved by matching the flow rate of buffered being transferred
into the recirculation tank to the flow rate of filtrate being
removed from the system. At the end of the diafiltration, the pool
was further concentrated in UF2 mode to a final concentration
set-point of greater than 170 g E25/L (35 L). This phase was also
performed at an elevated temperature set point of 45.degree.
C.+/-5.degree. C. Next, a low pressure drop mixing was performed
where the feed pump was controlled to maintain a 5-10 psig pressure
drop across the feed channel. A sample was pulled and a spec scan
was performed to confirm the concentration prior to recovery. The
concentration of this sample was 191 g E25/L and the pool volume
was 31.9 L. A graph of the process parameters over time were
comparable to those observed and summarized for the above FIG.
17.
TABLE-US-00018 TABLE 14 Normalized throughput Normalized Flux
Process Phase (g/ft.sup.2/hr) (LMH/psig) UF1 45.1 3.21 DF 25.9 2.51
UF2 121.4 0.79
[0147] Just prior to product recovery, a 30 mL sample was pulled
and analyzed for a titer of bioburden. The results of this test
were below the detection limit (<0.13 CFU/mL). The protein pool
was recovered by a series of steps. First the product was displaced
from the membrane in a single pass mode using 5 L of DF buffer
added to the retentate line. The product was filtered into a
recovery tank through a 7.4 ft.sup.2, 0.22 microns
sterilizing-grade guard filter followed by a 2 ft.sup.2, 0.22
microns sterilizing-grade final filter. The pool in the
recirculation tank was then pumped from the tank using the rotary
lobe feed pump. Next, the residual protein solution was displaced
from tank and feed line with a 5 psig nitrogen gas blow down. The
final phase was a blow down of the membrane unit, which contained
mostly DF buffer from the initial product displacement. This phase
also used 5 psig nitrogen gas applied to the highest point on the
retentate line. The recovered pool was diluted first to about 153 g
E25/L using DF buffer. Finally the pool was conditioned into the
final formulation of 20 mM histidine/200 mM arginine-HCl/0.04%
polysorbate 20/pH 6.0. The volumes of the recovered pool, diluted
pool, and conditioned pool were then analyzed for protein
concentration. Table 15 summarizes the results. Post process, the
membrane was regenerated as described above.
TABLE-US-00019 TABLE 15 Volume Concentration Yield or (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 1,082 5.5438 5,998.4 100 Recovered 34.95
167.08 5,839.8 97.4 Pool Diluted 38.2 152.14 5,810.3 96.7 Pool
Example 7
[0148] Effect of Elevated Temperature on Product Quality E25
samples at 30 g/L and 150 g/L in histidine and Q buffers were kept
a various temperatures for 24 hours. Samples were taken for
turbidity measurements and SEC assays. The results of turbidity
versus temperature for E25 at 30 g/L in Q buffer are shown in FIG.
18. FIG. 19 shows the amount of soluble aggregate of E25 (150 g/L
in 50 mM histidine buffer, pH 6.0) observed over time and at
temperatures of 23.degree. C., 40.degree. C., 50.degree. C.,
60.degree. C. and 70.degree. C. The four time intervals (time of 0
hours, 4 hours, 7.5 hours, and 24 hours) for each of these
temperatures is shown as the cluster of four bars from left to
right as 1810 and 1910, in FIGS. 18 and 19. The solution turbidity
was essentially unchanged after 24 hours at 60.degree. C. No
significant soluble aggregate of E25 was observed below 70.degree.
C. suggesting the product samples were substantially stable up to
at least 60.degree. C. and at least 24 hours.
Example 8
[0149] Effect of Elevated Temperature on Bioburden E25 samples at
30 g/L in both arginine and histidine buffers were inoculated with
10.sup.3 colony forming units per mL for two challenge organisms: a
Gram positive strain (Staphylococcus aureus); and one Gram negative
strain (Pseudomonas chlororaphis). Samples were taken after 1.5
hours and 6 hours. The results shown in the bar charts of FIGS. 20
and 21 indicate that these challenge organisms both decreased with
increasing temperature. The three temperature intervals (temp of
25.degree. C., 40.degree. C., and 50.degree. C. hours) for each
observed time interval is shown as the cluster of three bars from
left to right as 2010 and 2110, in FIGS. 20 and 21. The
inoculations shown were conducted in arginine buffer with protein
concentrations of 30 g/L.
Example 9
[0150] Effect of Elevated Temperature on Process Flux E25 samples
at 10 g/L in 0.2M arginine, 25 mM histidine, pH 6.0 buffer were
evaluated for their influence on flux versus transmembrane pressure
(TMP). FIG. 22 shows that raising the system temperature also
increased the process flux during the UF/DF operations. Flux
excursions at various bulk concentrations and three different
temperatures of 23.degree. C. (2210), 40.degree. C. (2220), and
46.degree. C. (2230) were performed. The mass transfer coefficient
and filtrate flux increased by about 2 to about 3 fold providing
considerably reduced process times.
Example 10
[0151] High Concentration Formulation of rhuMAB anti-CD20 ("2H7") A
pilot scale UF system was used to concentrate and formulate rhuMAb
anti-CD20 (2H7; a recombinant human monoclonal antibody). Example 1
was repeated with the following exceptions. The Millipore Pelicon
Ultrafiltration/Diafiltration systems were assembled with a
17.5-sqft, regenerated cellulose composite membrane, with a pore
size of 30,000-Dalton. The system consisted of a membrane holder, a
Viking S1 L rotary lobe feed pump, 1/2'' 316 L stainless steel
recirculation piping, and a 40-L recirculation vessel. Backpressure
regulating valves were H. D. Baumann, Inc. The temperature of the
glycol fed the tank jacket was regulated higher or lower as needed
using an electric heat exchanger, a cold glycol supply, or
both.
[0152] During this run, the feed flow rate was set to a constant
rate of 8.5 L/min (approximately 0.5 L/min/ft.sup.2). FIG. 23
displays the value trends over time for feed flow rate (210) scaled
from 0 to 20, pH (212) scaled from 2 to 12, filtrate flow rate
(250) scaled from 0 to 5, recycle tank level (2320) scaled from 0
to 45, and retentate dP (2350) scaled from 0 to 100 during the
various phases or mode of the process including UF1 (10), DF1 (20),
and UF2 (30).
[0153] During the ultrafiltration and diafiltration operations the
system used constant retentate pressure followed by a constant
feed/retentate delta pressure control scheme to control the flux
through the membrane. The diafiltration buffer (30 mM sodium
acetate/pH 4.9) was prepared in a separate tank. The temperature of
this buffer was set to 45.degree. C. prior to the process for
accurate temperature control through the entire process. Prior to
processing, the system storage solution (0.1 N NaOH) was flushed in
a single pass to drain mode first with 1 L/ft.sup.2 water for
injection (WFI) and then 1 L/ft.sup.2 diafiltration buffer. After
the flushes the system was equilibrated by recirculating 0.5
L/ft.sup.2 diafiltration buffer for 10 min. The pH of the
recirculated solution was checked to confirm the equilibration.
[0154] The protein pool resulting from a preceding Q-Sepharose
chromatography step was measured to be 2.31 g 2H7/L and had a
volume of 356 L. The protein was in a solution of 6 mM HEPES free
acid/19 mM HEPES sodium salt and 25 mM sodium acetate that had been
pH adjusted to 5.3 with 0.5 M acetic acid. Just prior to the run,
the temperature set point of this pool was set to 45.degree. C. To
begin the run the protein pool was transferred to the recirculation
vessel through a 0.22 microns sterilizing grade filter to a level
of about 40 L in the tank. In the vessel the pool was agitated via
the top mounted impeller and the temperature was maintained at
40-50.degree. C.
[0155] Because the incoming pool was larger then the recirculation
vessel, the UF1 process began in fed-batch mode (see FIG. 23). In
this mode, Q-Pool was added to the recirculation vessel at
approximately the same rate at as filtrate passes thought the TFF
membrane to drain. After the remaining Q-Pool has transferred to
the recirculation vessel, the UF1 process continued in batch mode.
During the UF1 the pool was concentrated to about 50 g 2H7/L (about
16 L). The pool was then diafiltered with 10 diavolumes of
diafiltration buffer. During diafiltration the temperature was
maintained between 40 and 50.degree. C. The diafiltration was
performed at a constant volume, which was achieved by matching the
flow rate of buffered being transferred into the recirculation tank
to the flow rate of filtrate being removed from the system. At the
end of the diafiltration, the pool was further concentrated in UF2
mode to a final concentration target set point of 190 g 2H7/L (4.3
L). See in FIG. 23 the incorporation of constant dP control at 50
psig at the end of this phase. This phase was also performed at an
elevated temperature set point of 45.degree. C.+/-5.degree. C.
Next, a low pressure drop mixing was performed where the feed pump
was controlled to maintain a 20 psig pressure drop across the feed
channel. A sample was pulled and a density measurement was
performed to confirm the concentration prior to recovery. The
concentration of this sample was 189 g 2H7/L. Table 16 summarizes
the throughput and flux results.
TABLE-US-00020 TABLE 16 Normalized throughput Normalized Flux
Process Phase (g/ft.sup.2/hr) (LMH/psig) UF1 32 4.8 DF 56 2.4 UF2
267 1.6
[0156] The protein pool was recovered by a series of steps. First
the product is displaced from the membrane in a single pass mode
using 0.2 L of DF buffer added to the retentate line. The product
is filtered into a recovery tank through a 0.22 microns
sterilizing-grade final filter. The pool in the recirculation tank
was then pumped from the tank using the rotary lobe feed pump.
Next, the residual protein solution is displaced from tank and feed
line with a 5 psig nitrogen gas blow down. The final phase was a
blow down of membrane unit, which now contains DF buffer from the
initial product displacement. This phase also used the 5 psig
nitrogen gas applied to the highest point on the retentate
line.
[0157] If necessary, the recovered pool was diluted first to about
175 g 2H7/L using dilution buffer (30 mM sodium acetate, pH 5.3).
Finally, the pool is diluted down to a target concentration of 150
g 2H7/L and conditioned into the final formulation of 30 mM sodium
acetate, 7% trehalose, 0.03% polysorbate 20, pH 5, via a 7.times.
conditioning buffer (30 mM sodium acetate, 49% trehalose, 0.21%
polysorbate 20, pH 5.3). The volumes of the recovered pool, diluted
pool, and conditioned pool were then analyzed for protein
concentration. Table 17 presents the results.
TABLE-US-00021 TABLE 17 Concentration Yield or Volume (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 355.81 2.31 821.92 100.0 Recovered 4.64
180.02 835.3 101.6 Pool Final Pool 4.871 149.40 727.7 88.5 Note:
Yields include loss due to sampling. Recovered pool volume and
concentration include addition of buffer displacement.
[0158] Post process, the membrane was regenerated using 0.1 N NaOH,
1 L/ft.sup.2 single pass flush followed by 0.5 L/ft.sup.2 total
recirculation for 30 min. This was followed by 1 L/ft.sup.2 PW
flush. This was followed by a total recirculation of 0.5 L/ft.sup.2
1.4% Minncare solution for 30 min. The system was again flushed
with 1 L/ft.sup.2 PW and finally recirculated for 15 min with 0.1 N
NaOH and stored.
Example 11
[0159] High Concentration Formulation of rhuMAb anti-CD20 A pilot
scale UF system was used to concentrate and formulate rhuMAb
anti-CD20 (2H7) for use in a human phase I clinical study in a GMP
manufacturing facility. Example 10 was repeated with the following
exceptions.
[0160] The protein pool resulting from a preceding Q-Sepharose
chromatography step was measured to be 3.729 g 2H7/L and had a
volume of 262 L. The protein was in a solution of 6 mM HEPES free
acid/19 mM HEPES sodium salt and 25 mM sodium acetate that had been
pH adjusted to 5.3 with 0.5 M acetic acid. Just prior to the run,
the temperature set point of this pool was set to 45.degree. C. To
begin the run the protein pool was transferred to the recirculation
vessel through a 0.22 microns sterilizing grade filter to a level
of about 40 L in the tank. In the vessel the pool was agitated via
the top mounted impeller and the temperature was maintained at
40-50.degree. C.
[0161] During the UF1 the pool was concentrated to about 50 g 2H7/L
(about 20 L). FIG. 24 displays the value trends over time for
recycle tank level (210) scaled from -0.713963 to 295.989,
retentate dP (2420) scaled from -0.237899 to 98.6629, feed flow
rate (250) scaled from -0.356981 to 147.994, and filtrate flow rate
(2450) scaled from -0.118994 to 49.3315 during the process. The
pool was then diafiltered with 10 diavolumes of diafiltration
buffer. During diafiltration the temperature was maintained between
40 and 50.degree. C. The diafiltration was performed at a constant
volume, which was achieved by matching the flow rate of buffer
being transferred into the recirculation tank to the flow rate of
filtrate being removed from the system. At the end of the
diafiltration, the pool was further concentrated in UF2 mode to a
final concentration target set point of 190 g 2H7/L (5.25 L). Note
in FIG. 24 the incorporation of constant dP to 40 psig control at
the end of this phase. This phase was also performed at an elevated
temperature set point of 45.degree. C.+/-5.degree. C. Next, a low
pressure drop mixing was performed where the feed pump is
controlled to maintain a 20 psig pressure drop across the feed
channel. A sample was pulled and a density measurement was
performed to confirm the concentration prior to recovery. The
concentration of this sample was 194 g 2H7/L. Table 18 summarizes
the throughput and flux results.
TABLE-US-00022 TABLE 18 Normalized Normalized throughput Flux
Process Phase (g/ft.sup.2/hr) (LMH/psig) UF1 51 3.8 DF 46 2.2 UF2
286 1.6
[0162] Just prior to product recovery, a 30 mL sample was pulled
and submitted for detection and titer of bioburden. The results
were negative (i.e., <0.13 CFU/mL). The protein pool was
recovered by the series of steps of Example 10. The volumes of the
recovered pool, diluted pool, and conditioned pool were then
analyzed for protein concentration. Table 19 presents the results.
The membrane was regenerated as in Example 10.
TABLE-US-00023 TABLE 19 Concentration Yield or Volume (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 262 3.72 977 100 Recovered 5.0 174.0
863.0 88.3 Pool Diluted 5.421 149.6 811.0 83.0 Pool
Example 12
[0163] High Concentration Formulation of rhuMAb anti-CD20 GMP
Example 11 was repeated with the following exceptions. The protein
pool resulting from a preceding Q-Sepharose chromatography step was
measured to be 5.106 g 2H7/L and had a volume of 196 L. The protein
was in a solution of 6 mM HEPES free acid/19 mM HEPES sodium salt
and 25 mM sodium acetate that had been pH adjusted to 5.3 with 0.5
M acetic acid. Just prior to the run, the temperature setpoint of
this pool was set to 45.degree. C. To begin the run the protein
pool was transferred to the recirculation vessel through a 0.22
microns sterilizing grade filter to a level of about 40 L in the
tank. In the vessel the pool was agitated via the top mounted
impeller and the temperature was maintained at 40-50.degree. C.
[0164] During the UF1 the pool was concentrated to about 50 g 2H7/L
(about 20 L). FIG. 25 displays the value trends over time for
recycle tank level (210) scaled from 0 to 300, retentate dP (2520)
scaled from 0 to 100, feed flow rate (250) scaled from 0 to 150,
and filtrate flow rate (2550) scaled from 0-50 during the process.
The pool was diafiltered with 10 diavolumes (10.times.) of
diafiltration buffer. During diafiltration the temperature was
maintained between 40 and 50.degree. C. The diafiltration was
performed at a constant volume that was achieved by matching the
flow rate of buffer being transferred into the recirculation tank
to the flow rate of filtrate being removed from the system. At the
end of the diafiltration, the pool was further concentrated in UF2
mode to a final concentration target setpoint of 190 g 2H7/L (5.26
L) again utilizing constant dP control at the very end of this
phase (see FIG. 25). This phase was also performed at an elevated
temperature set point of 45.degree. C.+/-5.degree. C. Next, a low
pressure drop mixing was performed where the feed pump is
controlled to maintain a 20 psig pressure drop across the feed
channel. A sample was pulled and a density measurement was
performed to confirm the concentration prior to recovery. The
concentration of this sample was 191 g 2H7/L. Table 20 summarizes
the throughput and flux results.
TABLE-US-00024 TABLE 20 Normalized Normalized throughput Flux
Process Phase (g/ft.sup.2/hr) (LMH/psig) UF1 67 3.6 DF 47 2.1 UF2
292 1.8
[0165] Just prior to product recovery, a 30 mL sample was pulled
and submitted for detection and titer of bioburden. The results
were negative (i.e., <0.13 CFU/mL). The protein pool was
recovered by a series of steps as in Example 11. The volumes of the
recovered pool, diluted pool, and conditioned pool were then
analyzed for protein concentration. Table 21 presents the results.
The membrane was regenerated as in Example 1 l.
TABLE-US-00025 TABLE 21 Concentration Yield or Volume (L) (g/L)
Mass (g) {Loss} (%) Q-Pool 196 5.106 1000 100 Recovered 4.9 187.1
918.0 91.8 Pool Diluted Pool 6.075 150.9 916.9 91.7
[0166] All publications, patents, and patent documents are
incorporated by reference herein in their entirety, as though
individually incorporated by reference. The disclosure has been
described with reference to various specific and preferred
embodiments and techniques. However, it should be understood that
many variations and modifications can be made while remaining
within the spirit and scope of the disclosure.
* * * * *