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\begin{document}

\title{Fully Integrated downstream process to enable Next Generation
Manufacturing}



\author[1]{Irina Ramos}%
\author[1]{Nikunj Sharda}%
\author[2]{Ramon Villafana}%
\author[3]{Kevin Hill-Byrne}%
\author[1]{Kang Cai}%
\author[3]{Joanna Pezzini}%
\author[1]{Jon Coffman}%
\affil[1]{AstraZeneca Medimmune}%
\affil[2]{Inhibrx LP}%
\affil[3]{PAK BioSolutions VA USA}%


\vspace{-1em}



  
  \date{June 16, 2024}


\begingroup
\let\center\flushleft
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\selectlanguage{english}
\begin{abstract}
Next generation manufacturing (NGM) has evolved over the past decade to
a point where large biopharmaceutical organizations are making large
investments in the technology and considering implementation in clinical
and commercial processes. There are many well-considered reasons to
implement NGM. For the most part, organizations will not fund NGM unless
the implementation benefits the funding organization by providing
reduced costs, reduced time or additional needed capabilities.
Productivity improvements gained from continuous purification are shown
in this work, which used a new system that fully integrates and
automates several downstream unit operations of a biopharmaceutical
process to provide flexibility and easy implementation of NGM. The
equipment and automation supporting NGM can be complicated and
expensive. Biopharmaceutical Process Development considered
two~options:~(1) design~its own NGM~system or (2) buy a pre-built
system.~PAK BioSolutions (Virginia, US),~provides a turn-key~automated
and integrated system that~can~operate up to four continuous
purification stages simultaneously, while maintaining a small footprint
in the manufacturing plant. The PAK system provides significant cost
benefits (\textasciitilde{}10x lower) compared to the alternative --
integration of many different pieces of equipment through a Distributed
Control System (DCS) that would require significant engineering time for
design, automation and integration. Integrated and Continuous
Biomanufacturing can lead to significant reductions in facility size,
reduced manufacturing costs, and enhanced product quality when compared
to~the traditional~batch mode of operation. The PAK~system~uses new
automation~strategies~that~robustly link~unit operations.~We present the
optimized process fit, sterility and bioburden control strategy, and
automation features (such as pH feedback control and in-line detergent
addition) that enabled continuous operation of a 14 day end-to-end
monoclonal antibody purification process at the clinical manufacturing
scale.%
\end{abstract}%



\sloppy


\section*{Introduction}

{\label{introduction}}

Continuous manufacturing processes have gained significant interest from
biopharmaceutical companies in recent years due to their potential to
provide large productivity increases over traditional batch production
modes (Konstantinov \& Cooney, 2014; Pollock, 2017; Arnold 2018). These
productivity increases allow a company to produce more biopharmaceutical
product from an existing manufacturing facility or build new facilities
with significantly reduced footprint compared to traditional designs.

Progress on upstream continuous processes has been widely published and
presented at industry conferences. Companies have reported titers for
perfusion processes reaching 2.3 g/L per day and beyond at scales over
500L (Coffman, 2021). This higher cell culture productivity results in a
purification bottleneck for traditional batch purification processes.
Perfusion cell culture volumes are typically on the order of 1.5 cell
culture vessel volume per day over 30 days and have viable cell
densities (VCD) and titers that vary over 10X across this duration. A
traditional 4 x 12,000L facility can now be replaced with a 5 x 2,000L
facility that outputs the same product mass, but at a cost of over 6X
more cell culture volume that must be forward processed (Arnold, 2018).
Purification facilities are not equipped with vessels to store the
entire batch volume and therefore, at a minimum, must process the
material through the purification unit operations in multiple
sub-batches.

Continuous purification unit operations have matured in recent years and
technologies now allow a full end-to-end continuous process without the
need for large hold vessels (Pezzini, 2018). These technologies
eliminate the challenge previously presented by large perfusion harvest
volumes. In a continuous purification process, each unit operation in
the process is connected in series and operated simultaneously. Since
the material is continuously forward processed from one step to the
next, there is no need to hold large intermediate volumes in tanks.
Furthermore, the purification process keeps pace with the perfusion
bioreactor and the full batch of drug substance is purified within 12
hours of completion of the cell culture operation (Coolbaugh, 2021).

Productivity improvements gained from continuous purification extend
beyond those achieved with perfusion cell culture processes (Klutz,
2015). Regardless of cell culture method, continuous purification offers
benefits of reduced equipment size and footprint, reduced capital and
consumable costs, and decreased run duration. Facility footprint is
significantly reduced by elimination of product hold tanks and the need
for utilities to clean and sterilize stainless steel piping systems. Low
flow rates allow for smaller piping diameters that can be provided by
single use tubing instead of stainless-steel piping. Advanced automation
allows processes to operate all steps 24/7 without intervention compared
to one step performed per shift, greatly improving utilization of
existing facility and personnel resources. Smaller chromatography
columns are cycled continuously throughout a run, reaching more than 35
cycles compared to 3 to 5 cycles on a larger column for a traditional
batch process. As a result, resin utilization is maximized and resin
costs are reduced (Otez, 2020). Although the capture chromatography step
sees the most consumable cost reduction in a continuous purification
process, labor and time savings for an end-to-end continuous process are
notable (Mahal, 2021).

End-to-end continuous purification processes have been published
previously by our group and others (Godawat, 2015; Pezzini, 2018,
Coolbaugh, 2021). They include processes up to 25 days in duration and a
demonstration of bioburden control. Even so, significant opportunities
for process maturity exist for certain unit operations, equipment
hardware and automation. Here, we present the latest generation of
continuous purification processes with an off-the-shelf Pilot PAK
continuous purification system (PAK BioSolutions, Virginia, US) that
includes single-use GMP flow kits, advanced control strategies, and
supervisory control and data acquisition (SCADA).

In this work, we demonstrate a 14-day continuous monoclonal antibody
(mAb) purification process for a 50L perfusion bioreactor. Bioburden was
controlled within specifications for the duration of the process. A
single PAK automated system with four flow kits installed was used to
perform the following steps; harvested perfusion material was processed
through a Protein A capture chromatography step, a low pH virus
inactivation step, a depth filter, a 0.2\selectlanguage{greek}µ\selectlanguage{english}m sterile filter and an anion
exchange membrane. The process performed did not require a cation
exchange chromatography step to meet process impurity specifications.
Ultrafiltration/ diafiltration operations were performed separately
after breaking down and setting up the same PAK system in a UFDF setup.

As demonstrated in this work, continuous cell culture and purification
processes have advanced beyond proof-of-concept studies. The significant
advancements in off-the-shelf equipment now enable routine operation of
continuous purification processes in pilot and GMP facilities. This
equipment reduces footprint requirements for an end-to-end process by
75\% and offers a level of automation that reduces hands on time by 90\%
compared to batch. Furthermore, we present low pH viral clearance data
required to implement a continuous purification process for a GMP
facility.

\section*{Materials}

{\label{materials}}

\subsection*{Equipment and Consumables}

{\label{equipment-and-consumables}}

A Pilot PAK end-to-end continuous purification system from PAK
BioSolutions (Virginia, US) was used to perform all purification process
steps. Along with the Pilot PAK system were four Pilot PAK flow kits
from PAK BioSolutions (Virginia, US) that were used to link unit
operations together. MabSelect PrismA\selectlanguage{ngerman}® and Sephadex G-25 Coarse® size
exclusion chromatographic (SEC) resins were purchased from Cytiva
(Massachusetts, US). The three chromatography column housing units were
Vantage L Laboratory Column VL 44 x 250 from MilliporeSigma
(Massachusetts, US). The Opticap Sterile 0.2\selectlanguage{greek}µ\selectlanguage{english}m filter and the Mustang Q
Membrane Capsule were purchased from MilliporeSigma (Massachusetts, US)
and Pall (New York, US) respectively. Tryptic Soy Agar plates were
purchased from VWR (Pennsylvania, US).

\subsection*{Buffers and Solutions}

{\label{buffers-and-solutions}}

The buffer system utilized for dual-column capture chromatography
(MabSelect PrismA\selectlanguage{ngerman}®) included six buffers: (1) 50 mM Tris, pH 7.4; (2)
50mM Tris, 0.5 M NaCl, pH 7.4; (3) 25mM sodium acetate, pH 3.6; (4)
100mM acetic acid (5) 1N NaOH and (6) 2\% benzyl alcohol, 100 mM sodium
acetate, pH 5.0. For the Viral Inactivation (VI) step, the solution
system utilized 1M acetic acid and 1M tris base. All buffers were tested
for bioburden and only qualified for use with negative results.

\subsection*{Protein}

{\label{protein}}

The monoclonal antibody (mAb1) used in the continuous process studies
was a humanized IgG1 produced from Chinese hamster ovary (CHO) cells by
AstraZeneca. This protein has a molecular weight of 146 kDa, as
determined experimentally by mass spectroscopy and has a theoretical
isoelectric point of 7.15, as determined by amino acid sequencing.

The monoclonal antibody (mAb2) used in viral clearance studies was a
humanized IgG1 produced from Chinese hamster ovary (CHO) cells by
AstraZeneca. The mAb2 intermediate (17 mg/mL, pH 4.5) purified by
protein A chromatography was used in the virus inactivation study as
starting material.

\section*{Methods}

{\label{methods}}

\subsection*{Process overview}

{\label{process-overview}}

The purification process for this study included an automated and
integrated dual-column chromatography step (DCC) (Angarita, 2015),
leading into a continuous virus inactivation (VI) step and a final
flow-through polishing AEX step. Upstream and downstream of the DCC step
were 10L glass vessels to accommodate purification of 50L of cell
culture material per day. The process flow path, the connections between
each unit operation in the process and the automation to control and
acquire the data from all unit operations in parallel, were executed
with four 500mL flow kits and the Pilot PAK system from PAK BioSolutions
(Virginia, US). The flow path of the process and the PAK flow kits are
shown in Figure 1 and Figure 2, respectively. The 50L per day of
harvested material was combined from several different perfusion
bioreactors that had run prior to this study. The bags were chosen and
connected in order to simulate the titer profile of a typical perfusion
bioreactor process.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image1/image1}

\end{center}
\end{figure}

Figure 1. Process flow path in PAK System.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image2/image2}

\end{center}
\end{figure}

Figure 2. Flow path of the PAK single-use flow kit for individual
process steps (chromatography, virus inactivation, membrane,
filtration). Includes pH, conductivity (Cond), and pressure (P)
instrumentation.

The 500mL flow kits are 3D-printed and include sensors (pH,
conductivity, UV280nm, UV300nm, pressure), in-line mixers for buffer
addition (pH titration, conductivity adjustment, in-line dilution), a
vented mixing vessel and a sample port (Figure 2).

\subsection*{Sterilization and Sanitization
Pre-Run}

{\label{sterilization-and-sanitization-pre-run}}

Bioburden control of the continuous purification process required a
combination of sanitization and sterilization methods. The Pilot PAK
flow kits and associated tubing assemblies were provided
gamma-irradiated. Chromatography columns (MabSelect PrismA\selectlanguage{ngerman}® used for
capture step and SEC column used for VI step) were sanitized with 1N
sodium hydroxide during automated start-up sequences, along with the
associated flow paths that included pH, conductivity and UV
instrumentation. The 0.2\selectlanguage{greek}µ\selectlanguage{english}m sterile filter and the Mustang Q membrane
were sanitized offline per the vendor specifications with 1N sodium
hydroxide prior to use. Glass vessels used as product collection tanks,
situated after the perfusion bioreactor and capture step, were
autoclaved with tubing. Consumable components such as bags and tubing
were gamma irradiated or autoclaved. Sanitary quick connects were used
to make any connections that were not sanitized prior to use.

\subsection*{Cell Culture}

{\label{cell-culture}}

Several 3L bioreactors were used to generate recombinant IgG via a
Chinese Hamster Ovary (CHO) cell line. These individual 3L bioreactors
were harvested using coupled Tangential Flow Filtration (TFF), from
which they were combined into 50 and 100L bags according to their
titers. The titer profile for the 14-day study presented here was
generated by ordering the bags based on their titer to represent a
typical 50L perfusion bioreactor process. These bags were then welded to
the flow path and pumped at a rate of 50L/day into a 10L glass vessel
before the dual-column chromatography step. The titer of the material
used ranged from 0.2 -- 2.34 g/L and totaled 490g of harvested mAb.

\subsection*{Dual Column Chromatography for Protein A
Capture}

{\label{dual-column-chromatography-for-protein-a-capture}}

The target molecule was captured with two 4.4 cm diameter x 7.7 cm
height chromatography columns using MabSelect PrismA\selectlanguage{ngerman}® resin at a
capacity of 60 g/L. The columns for the capture step were operated using
a dual-column strategy which operated in a typical bind and elute
methodology, enabling a continuous load and discontinuous product
elution. The size of the columns is optimized for protein mass and
process time, resulting in higher resin utilization. Parallel column
operation maintains one of the two columns in the load phase at all
times. While one column is in the load phase, the other column cycles
through the process equilibration, wash, elution, strip and sanitization
buffers (Figure 3). The columns were loaded to 60 g/L every cycle; the
load duration and frequency of elution varied with the change in titer
(load volume was manually input into the PAK automation control screen
daily). The elution collection criteria was specified as 5 column
volumes (CVs), and collection began after the completion of the first
CV. The eluate was collected in a 10L glass vessel downstream of the two
columns.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image3/image3}

\end{center}
\end{figure}

Figure 3. Diagram illustrating the dual column operation. Condition
indicates pre elution phases such as equilibration and wash. Regen
indicates the combination of the post-elution phases: strip,
sanitization and re-equilibration. DBC stands for dynamic binding
capacity. Grey shade illustrates protein loaded into the column.

\subsection*{Virus Inactivation}

{\label{virus-inactivation}}

The viral inactivation strategy implemented in this study used a plug
flow reactor made from a 4.4 cm diameter x 20 cm height SEC (Size
Exclusion Chromatography) column with Sephadex G-25 Coarse\selectlanguage{ngerman}® resin to
achieve a target residence time of 15 minutes. Acidification was
performed before the SEC column and neutralization was performed post
column. The capture product was titrated to an acidification target of
pH 3.5 with an allowable range of ± 0.1. Acid was added in-line with the
product line and the acidified product immediately flowed through the
flow kit's static mixer and pH probe. A feedback PID controller is used
to continuously adjust the acid pump flow rate to ensure the pH is
within the desired range. After the pH is within the acidification
target range for an appropriate amount of time, the product is directed
through the SEC column which was equilibrated at pH 3.5. Post SEC
column, an in-line addition of base is continuously fed to neutralize
the acidic product to a target of pH 7.2 with an allowable range of ±
0.2. The viral inactivation procedure during this process began on day 8
of the 14-day run. The start day for this step is determined based on
(1) product quality impact from hold time at capture product conditions
and (2) system flow rate lower limit to allow VI to start and run for
extended duration. For this study VI was run in three sub-batches. For a
future optimized process, we aim to operate without pausing.

\subsection*{Anion Exchange
Chromatography}

{\label{anion-exchange-chromatography}}

The product from VI step flowed through a filter train comprised of two
Opticap® Sterile 0.2\selectlanguage{greek}µ\selectlanguage{english}m filters on either side of a Mustang Q\selectlanguage{ngerman}® Membrane
Capsule. This filter train was prepped offline, prior to its connection
to the system according to the vendor specifications. This prep work was
done the day before the filter train was connected to the system (day
7). The filter train was connected on day 8; this timeline coincided
with the start of the viral inactivation step. From here, material was
collected in a final 50L product collection bag. This final polishing
chromatography step was performed with a single-use membrane in a
flowthrough mode. The membrane was not reused. The 0.2\selectlanguage{greek}µ\selectlanguage{english}m filter
downstream of the Mustang Q membrane ensured that no potential bioburden
reached the final product collection bag.

\subsection*{Bioburden Sampling and Testing
Strategy}

{\label{bioburden-sampling-and-testing-strategy}}

Samples were taken from four points in the process: (1) the cell culture
harvest 10L vessel, (2) Protein A product 10L vessel, (3) VI product
flow kit vessel and (4) final product collection bag. Samples were taken
from sample ports that were specifically installed during set-up and
were clamped closed when not in use to prevent potential exposure to the
environment. Samples were taken aseptically using 10 mL syringes, about
24 hours apart. Samples were aliquoted into individual, labelled tubes
under a biosafety cabinet in order to preserve sterility. Once in the
tubes, the samples were stored at 2-8\selectlanguage{ngerman}°C until they were ready to be
submitted for analysis (see section Analytical Testing

for more details). This sampling method was repeated every day of the
run, for each sampling point. The protein A bindable testing was done in
order to record the daily titer at each stage. The HCP and DNA
concentration assays tested the impurity level at each stage and
therefore see how effective each unit operation was at removing the
impurities. The SE-HPLC testing was done to determine the monomer purity
at each stage and therefore determine daily product quality across the
process. The samples for testing bioburden were taken and plated onto
culture plates under a biosafety cabinet using L-shaped spreaders. Once
done, the plates were stored inside of an incubation chamber running at
35 °C. The plates were monitored every day until 72 hours had passed in
order to ensure the sterility of the system. These sampling and
bioburden plating methods were repeated every day for each sampling
point, through the end of the run.

\subsection*{Analytical Testing}

{\label{analytical-testing}}

Size exclusion-high performance liquid chromatography (SE-HPLC)

{\label{size-exclusion-high-performance-liquid-chromatography-se-hplc}}

SE-HPLC was performed using a Waters 515 pump, a Waters 2487 Dual
Absorbance Detector (Waters) and a Rheodyme 77,251 injector and a TSK
Gel G3000 SWXL column (300 mm × 7.8 mm, Tosoh Biosep). The volume of
injection was adjusted according to each sample concentration in order
to inject 250 \selectlanguage{greek}µ\selectlanguage{english}g of mAb1, and separation was performed at a flow rate of
1.0 mL/min. The running buffer was composed of 100mM sodium sulfate,
100mM sodium phosphate dibasic, pH 6.8. UV detection was performed at
280 nm. The results were reported as monomer peak percentage and high
molecular weight (HMW) content percentage of the total integrated area.

CHO DNA Real-Time PCR Assay

{\label{cho-dna-real-time-pcr-assay}}

The assay is performed in a 96-well plate format using a forward and
reverse primer set specific for the Chinese hamster short interspersed
repetitive element (SINE) DNA sequence. A unique 64 base-pair product is
generated from CHO genomic DNA calibrators and test samples within 40
cycles of the PCR. DNA is extracted from test samples and an extracted
control using the DNA Extractor Kit\selectlanguage{ngerman}® from Wako. Test samples, CHO
genomic DNA calibrators and controls are added to a reagent master mix
containing primers and SYBR Green. Cycling is performed in three steps
as follows: (1) 10 minutes at 95°C for polymerase activation; (2) 40
cycles of 15 seconds at 95°C for denaturation; (3) 40 cycles of 60
seconds at 60°C for annealing of primers and extension. Quantities of
DNA in the samples above LLOQ are regressed from an assay specific
standard curve generated by the Sequence Detection Systems (SDS)
software.

Gyros CHO Host Cell Protein (HCP) Assay

{\label{gyros-cho-host-cell-protein-hcp-assay}}

Standard curve, samples and controls are diluted in a proprietary Gyros
buffer and loaded into sample microplates. In-house prepared
biotinylated sheep anti-CHO HCP capture antibody and fluorescent labeled
sheep anti-CHO HCP detection antibodies are diluted and loaded into a
reagent microplate. The microplates are loaded on the Gyrolab instrument
and automated execution of the assay is initiated. Once all reactions
are complete in the CDs are read by the instrument using a fluorescence
detector at 647 nm. HCP quantities are regressed from the standard curve
generated by the software and corrected for dilution factor. Reported
sample concentrations are the mean concentration of the dilutions
yielding spike recovery values of 75-125\%.

HPLC Protein A bindable Assay for mAb concentration

{\label{hplc-protein-a-bindable-assay-for-mab-concentration}}

Protein A bindable HPLC is used to determine IgG concentration utilizing
Protein A affinity chromatography separation technique. Test sample is
injected onto a PA Immuno-Detection Sensor Cartridge (2.1 mmD x 30 mmL)
with PBS buffer as a mobile phase.~ The sample flows through the
cartridge where the target IgG specifically binds to the Protein A
immobilized on the cartridge and non-target components are washed out.~
The bound IgG is eluted with 0.1\% phosphoric acid in PBS, and detected
using UV absorbance at 280 nm.~ The IgG content is then quantified using
linear regression analysis of the standard curve prepared from the
corresponding calibration standards.~ Results are reported in \selectlanguage{greek}µ\selectlanguage{english}g/mL.

\section*{Results and Discussion}

{\label{results-and-discussion}}

This study successfully demonstrated a next-generation continuous
purification process, running for 14-days at a rate of 50L of cell
culture harvest material per day. Bioburden control was maintained for
the duration of the run and product quality results were within
acceptable limits. Advanced automation enabled tight control of the
process to meet specifications. Automation control strategies also
enabled break vessel volumes between 150 ml and 10L (20-fold or greater
reduction compared to batch). Facility footprint of the process was
reduced by 75\% compared to the batch process, primarily due to the
reduction in vessel volumes and ability to perform all three process
steps in a single system. Setup time was to 1.5 days and was followed by
14 days of hands-free operation with the exception of daily sampling
(90\% reduction in operator hours). These outcomes demonstrate the
robustness of this commercially available automation system for
continuous purification processes.

\subsection*{Bioburden Control}

{\label{bioburden-control}}

Most downstream operations for monoclonal antibody are not considered
aseptic and operate as low-bioburden. The guidance commonly used sets a
limit at 10 colony-forming units (CFU) per mL.

Bioburden control was demonstrated for the duration of the 14-day run.
Samples were taken daily from the harvest break vessel, protein A
product vessel, VI neutralized product vessel and AEX final product.
None of the sampling points tested positive for bioburden except, on day
7, when 1CFU/mL was observed for the protein A product vessel (below the
acceptable criteria of \textless{}10CFU/mL). This hit was attributed to
intentionally opening the protein A product vessel to fix an incorrectly
installed tubing assembly on the vessel. The bioburden objectives for
this process were clearly met. The 1CFU/mL sample from the protein A
elution product vessel was cleared by implementing a bioburden
mitigation strategy. This strategy consisted of six important steps: (1)
pausing the protein A step, (2) removing the contaminated vessel from
the system, (3) autoclaving a new vessel, (4) filtering material into
the new vessel under aseptic conditions, (5) integrating the new vessel
into the process and (6) flushing lines with sanitization buffer before
restarting the protein A step. No bioburden was detected after the
implementation of this strategy. Successful implementation of this
mitigation strategy demonstrates the ability to clear bioburden from an
ongoing process without significant process disruption, a necessity to
save a batch when operating a continuous operation.

\subsubsection*{Capture}

{\label{capture}}

Seventy protein A capture chromatography cycles were successfully run
during the 14-day process. The protein A capture step began
automatically once the harvest vessel reached a setpoint volume of 2L.
The capture step operated for the duration of the run.

The capture step was operated under a constant mass load principle,
targeting the resin capacity of 60g of mAb per L of resin: the load
volume onto the column changed as the titer varied throughout the run to
ensure that the grams of mAb loaded was the same for each cycle, as
illustrated in Figure 4. Load volume decreases as titers increase until
the middle of the run. Load volume then increases again as the titer
drops towards the end of the run. This strategy ensures a constant
concentration downstream of the protein A capture step, which greatly
simplifies the downstream process dynamics. The standard deviation for
the Protein A concentration over the duration of the run was 1.4 g/L.
With constant downstream concentration, the virus inactivation step does
not need to account for large mAb concentration changes that can affect
the titration profile. Additionally, polishing chromatography step
throughput can be calculated from volumetric throughput and the known
concentration. The capture step was triggered to stop once the harvest
vessel dropped to a low level due to lack of flow from the cell culture
feed.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image4/image4}

\end{center}
\end{figure}

Figure 4. Protein A load volume and titer for the duration of the run.
Black circles and triangles represent the start of the cycles for column
1 and 2, respectively.

Chromatogram overlays of the 35 cycles performed on each of the two
capture columns were generated for analysis (data not shown). Constant
peak height and position are observed across all cycles. Tailing was
noticed on some cycles, especially towards the end of the run. Analysis
of conductivity traces showed a change in transition curve profile for
these cycles which indicates a change in bed structure rather than a
change in impurity profile or issue with the UV meter. The root cause
was likely air on the columns. This can easily be mitigated for future
processes by installing a bubble trap prior to the protein A step, which
was not in place during this run.

\subsection*{Virus Inactivation}

{\label{virus-inactivation-1}}

The continuous low pH virus inactivation step was successfully
controlled to meet pH and residence time specifications. The virus
inactivation step was performed in 3 parts by starting and pausing this
step 3 times during the 14-day process. This pausing was required to
account for the change in rate of product elutions from the upstream
protein A step and maintain a constant residence time across the virus
inactivation column. The virus inactivation step was triggered to
automatically start when the protein A product vessel of 10L filled to a
high target volume. During the run, the virus inactivation step drained
the vessel until it reached a low target volume, at which point the
virus inactivation step automatically paused until the high target level
was reached again. Each of the three virus inactivation sub-runs lasted
approximately 1 day.

The virus inactivation step used an SEC column to provide the required
inactivation residence time. The SEC column models a plug flow reactor
and this choice benefits from the abundance of industry knowledge around
column packing and qualification. Additionally, residence time
distribution is independent of flow rate in a packed bed with no
liquid-solid mass transfer, unlike tubular flow reactor designs which
rely on dean vortices for mixing (Amarikwa, 2019). Other options, such
as alternating tank VI were not considered due to the increased space,
equipment, and automation requirements.

The pH of the capture product was titrated inline down to a pH of 3.5
with 1 M of acetic acid, before passing through the SEC column. After
exiting the SEC column, pH was titrated inline up to a target pH of 7.2
with 1M tris base. Inline pH adjustments were performed using feedback
control. The pH was well controlled for and was mostly maintained within
the specified ranges of \selectlanguage{ngerman}±0.1 and ±0.2 for the acidification and
neutralization, respectively. Any disturbances outside of the specified
ranges were diverted to waste until pH returned within spec and remained
stable for 2 minutes. pH data can be seen in Figure 5.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image5/image5}

\end{center}
\end{figure}\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image6/image6}

\end{center}
\end{figure}

c\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image7/image7}

\end{center}
\end{figure}

Figure 5. Virus inactivation pH trends. (a) Valve switch to in-line with
process during startup. (b) pH trends for sub-run 3. NOR is normal
operating range defined by the process description. (c) Virus
inactivation residence time trend during sub-run 3.

Residence time through the SEC column was also controlled to a setpoint
through feedback control. Rather than simply setting the pump to a fixed
flow rate to achieve the required residence time, the last pump in the
process changed flow rate to drive the overall mass flow rate of the
process to match the flow rate required to achieve the target residence
time in the SEC column. This strategy is required when running a truly
integrated process as fixed flow rates on any step will inevitably
result in vessel over/underflow events due to imperfectly calibrated
pumps.

\subsection*{Polishing}

{\label{polishing}}

A flow-through anion exchange membrane polishing step was performed for
this run. A second polishing step was not required to meet product
quality specifications. A 0.2\selectlanguage{greek}µ\selectlanguage{english}m filter was installed before and after
the Pall Mustang Q AEX membrane, which was immediately prior to the
final collection bag. The filter train, which consisted of the AEX
membrane and two 0.2\selectlanguage{greek}µ\selectlanguage{english}m filters (pre and post membrane), was swapped with
a fresh filter train once the capacity on either the filter or membrane
had been reached. The filters were sized in order to match their
capacity with the AEX membrane and were replaced at the same time. The
throughput was automatically determined by the PAK system, which alerted
the user when a swap was required.

The filter train was swapped once during the process, although ideally,
two swaps would have been performed. Unfortunately, the COVID-pandemic
impacted global supplies and limited the availability of single-use
materials like filters and membranes, thus only one swap was able to be
performed. This resulted in an intentional overloading of the second
Mustang Q to prevent cutting the run short. The first train was loaded
to 14.2 liters while the second was loaded to 30.4 liters. Aside from
the consequence of limited membrane area, the polishing step ran as
expected.

\subsection*{Product Quality}

{\label{product-quality}}

The continuous purification process performed as expected and
demonstrated acceptable product quality results. DNA, HCP, and monomer
purity measurements were performed. Samples were taken from the harvest
vessel, protein A product vessel, virus inactivation product flow kit,
and the final collection bag. The final collection bag contained all the
material collected from a single filter train. Product was diverted to a
new collection bag when the filter train was swapped. The final AEX
product was within acceptable ranges for percent monomer and host cell
protein on both days measured. DNA levels were acceptable on day 9 but
out of range on day 14. This was attributed to the intentional
overloading of the second AEX membrane as described above. Overall,
product quality was comparable to the batch downstream process, with the
exception of the day 14 AEX DNA datapoint. Results can be seen in Figure
6.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image8/image8}

\end{center}
\end{figure}\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image9/image9}

\end{center}
\end{figure}\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image10/image10}

\end{center}
\end{figure}

Figure 6. Product quality results from each sampling point across the
duration of the run. Black dashed line indicated the acceptable cut-off
for the final AEX product. Each datapoint is from a different day in the
process. AEX final product sample points are labeled with day of sample
collection. (a) HCP titers in ng/mL (acceptable cut-off below line). (b)
Percent monomer (acceptable cut-off above line). (c) DNA titers in ng/mL
(acceptable cut-off below line). Circled DNA datapoint on AEX step shows
out of spec sample due to intentional overloading of AEX membrane.

\subsection*{Advanced Automation Control
Strategies}

{\label{advanced-automation-control-strategies}}

A key objective of this work was to evaluate the ease of integrating
multiple unit operations, which requires complex automation strategies.
Unit operation integration was handled completely by the PAK system,
which removed the burden of in-house automation development and
implementation. With the PAK system, unit operation integration is
mediated by system-monitored break vessels. For this run, this included
the (1) harvest break vessel, (2) protein A capture product vessel, and
(3) VI product flow kit vessel. The harvest break vessel and the virus
inactivation neutralized product break vessel were controlled to a level
setpoint using PID feedback control over the vessel outlet and inlet
pump, respectively. The protein A product vessel used minimum, low,
high, and maximum level criteria to trigger downstream pause/restart
events. Overall, all vessels were controlled as intended. The
significance of this control strategy is that it eliminates the concern
of over/under flow of intermediate break vessels.

The protein A elution vessel changeover performed on day 9, as described
in the Bioburden Control section, resulted in a buildup of cell culture
material in the harvest vessel, as the system was paused during this
time while perfusion material continued to accumulate. The 10L capacity
of the harvest vessel coupled with its 50L overflow bag provided ample
time for an operator to swap vessels and restart the system without
further incident. The protein A product vessel triggered the downstream
operations to begin once reaching a target level of 7.5 liters. After
draining to a level of 1.5 liters, the downstream unit operations were
triggered to pause to allow the vessel to refill. These trigger levels
were manually changed through the HMI throughout the run to allow for
the trigger events to coincide with an operator onsite. This was due to
an abundance of caution and is not required for the process to operate.

A critical, yet often overlooked parameter in a continuous process is
the mass flow rate of the overall process. Rather than set each unit
operation to a specified flow rate and over-size break vessels to absorb
differences, the PAK system actively controls the flow rate of each
individual step and the system as a whole. This is achieved by
controlling the last product pump in series to a flow rate set point,
while regulating flow of upstream product pumps to maintain intermediate
break vessel level (the protein A capture step is an exception). As a
result of the vessel level control, any change in flow rate of the last
product pump results in a corresponding change in flow rate of the
upstream product pumps. This control strategy allows for low break
vessel volumes (150 ml for this 50 L cell culture per day process).

The mass flow rate set point (set for the last product pump in series)
is determined by the most critical process parameter. For this run, the
most critical process parameter was determined to be the virus
inactivation residence time. Therefore, the mass flow rate of the entire
process was controlled to ensure the VI residence time setpoint was met.
A schematic of this process is displayed in Figure 7.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image11/image11}

\end{center}
\end{figure}

Figure 7. Schematic of mass flow rate control. (1)
F\textsubscript{proAeluate}: protein A product material flow rate,
F\textsubscript{acid}: acid flow rate, F: total flow into VI SEC column,
F\textsubscript{sp}: set point flow rate, E\textsubscript{flow}: error
in target flow into SEC column. (2) E\textsubscript{level}: Error in
target vessel level, V: vessel level, V\textsubscript{sp}: set point
vessel level. (3) Ratio: target ratio of acid flowrate to product
flowrate.

The following steps are performed to adjust the pump flow rates to
ensure the virus inactivation residence time set point is met. First,
the total flow into the virus inactivation SEC column is calculated as
the sum of the protein A product feed (VI product pump) and the VI acid
pump flow rates. The flow rate set point that corresponds to the virus
inactivation residence time target is then subtracted from the total
inlet flow, providing the flow rate error. The flow rate error is fed
into a PID controller that controls the AEX product pump (the last
product pump in series in this process). Because this pump pulls from
the VI break vessel, any change in this flow rate will immediately
result in a change in VI product vessel level. The error of this vessel
is fed into a PID controller that controls the VI product pump to
maintain a constant VI vessel level. Finally, the VI acid pump output
increases or decreases so that the target ratio between the VI acid and
VI product pumps is always achieved.

The virus inactivation residence time for the third VI sub-run is
displayed in Figure 5c.

This series of independent control loops, connected only by the physical
properties of the VI product vessel allows for cohesive mass flow
control across unit operations. Additional unit operations can be added
between the VI and AEX steps with this strategy. In this case, flow
control cascades up through multiple break vessels.

The perfusion process required a break in mass flow rate at some point
in the process. The perfusion rate must be set constant and for this run
the VI residence time was controlled to a setpoint. Regardless of the
accuracy of pumps, there will inevitably be a difference in mass flow
rates between these two unit operations. The break in mass flow rate
occurred between the protein A chromatography step and the virus
inactivation step. The difference in mass flow rate was absorbed by the
protein A product break vessel. This is the reason the protein A product
vessel operated under min, low, high, max control rather than continuous
PID control and why the virus inactivation step ran as three separate
sub batches during this run. In future runs the VI step will operate
within a range of acceptable residence times which will allow VI flow
rates to adjust to match the changing output of the protein A eluate
vessel. This will enable this step to run continuously for the duration
of the process instead of periodically stopping and restarting.

\subsection*{Viral Clearance Data from continuous Virus
Inactivation}

{\label{viral-clearance-data-from-continuous-virus-inactivation}}

To ensure viral clearance across the flow through virus inactivation SEC
column, live virus spiking studies were performed. These studies aimed
to (1) characterize the breakthrough profile of virus and (2) to ensure
inactivation kinetics occur in the same manner to batch inactivation.

Before testing the column with the virus-spike, the breakthrough profile
of the column was characterized with blue dextran as a virus surrogate.
Blue dextran is a large molecule with a molecular weight of
\textasciitilde{}2000kDa (\textasciitilde{}10-20nm diameter). It is
commonly used to determine column void volume and is expected to behave
similarly to enveloped viruses, which are typically larger than 50 nm
(ICH Q5A, Food and Drug Administration, 1999). Blue dextran spiking
studies were performed at various flow rates across the SEC column to
characterize the breakthrough profiles of different column residence
times. Data can be seen in Figure 8.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image12/image12}

\end{center}
\end{figure}

Figure 8. Peak overlays of all blue dextran elutions. Variance in peak
height is due to variations in blue dextran spiking concentration and is
irrelevant to the results.

As can be seen from the results, breakthrough volume was independent of
flow rate, with a standard deviation of only 0.004 column volumes. This
can be explained by the Van Deemter Equation:

\begin{equation}
HETP=A+\frac{B}{u}+(C_{s}+C_{m})\bullet u\nonumber \\
 \end{equation}

Where A is the constant eddy diffusion parameter, B is the diffusion
coefficient in the longitudinal direction, and Cs and Cm are the
resistance to mass transfer in the stationary and mobile phases,
respectively. Because there is no mass transfer of blue dextran from the
liquid phase to the solid phase or vice versa, this portion of the
equation can be ignored. The flow regime is also operated at a
relatively large u, making the longitudinal diffusion part of the
equation negligible as well. This eliminates any dependence of velocity
on peak broadening. These results are significant because they
demonstrate a consistent breakthrough profile regardless of flow rate,
which may need to be adjusted throughout the course of a continuous
process. Therefore, breakthrough time across the column can be easily
calculated from the current flow rate, column volume, and void volume
coefficient. This is not necessarily true of other methods of flow
through virus inactivation that rely on Dean vortices for axial mixing.

In order to demonstrate that live virus behaves similar to blue dextran,
the column was spiked with inactivated xMuLV virus while running at a
residence time of 30 minutes. Three 6 mL fractions of eluate were
collected corresponding to before, during, and after the expected
breakthrough of virus. Results can be seen in Figure 9.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image13/image13}

\end{center}
\end{figure}

Figure 9. Breakthrough of XMuLV overlayed with a blue dextran run. Both
datasets performed at 7.3cm/hr.

The pre breakthrough fraction contained 2.85x10\textsuperscript{1} virus
copies/mL indicating little to no breakthrough as expected. The expected
breakthrough fraction contained 6.78x10\textsuperscript{6} virus
particles and the post breakthrough fraction contained
1.65x10\textsuperscript{6} virus particles. This follows a similar
profile to the blue dextran elution profile, and most importantly shows
virus breakthrough does not occur earlier than blue dextran.

In addition to breakthrough characterization, virus inactivation was
also evaluated at multiple residence times across the column. Virus
inactivation was measured at residence times of 2, 10, 20, and 30
minutes. Log inactivation data is shown in Figure 10.\selectlanguage{english}
\begin{figure}[H]
\begin{center}
\includegraphics[width=0.70\columnwidth]{figures/image14/image14}

\end{center}
\end{figure}

Figure 10. Log reduction values for both continuous and traditional
batch modes of low pH virus inactivation. Both datasets performed at pH
3.6 with the same molecule. Dots on continuous dataset indicate start of
collection of eluate sample while lines indicate the span of collection
of the sample.

The log inactivation data shows a similar trend to the batch
inactivation data for the same molecule under the same conditions. This
clearly shows there is no difference in the kinetics of inactivation
between batch and continuous methods.

\section*{Conclusion}

{\label{conclusion}}

We have successfully demonstrated a next generation continuous
purification process for a 14-day operation with bioburden control. The
PAK automated system for continuous purification provided the automation
control, automation hardware and single-use assemblies required.
Advanced control strategies controlled flow from one step to the next
and the mass flow rate of the entire process. The control parameters
were continuously monitored and set to alarm when limits were reached.
No additional automation control or SCADA system was required to perform
this work. Automation used was 21CFR Part 11 compliant and all product
contact materials met or exceeded USP Class VI specifications. The Low
pH continuous viral inactivation method used in this process was
supported by off-line virus spiking studies and results are comparable
to the batch process. Batch chromatography studies supported the dual
column operations performed in the continuous process. Additional
studies are required to confirm viral filtration efficacy at the reduced
flow rates used in the continuous process. The technology presented
enables implementation of continuous fully integrated processes at
scales up to 200L of cell culture material per day, which is a
significant step towards good manufacturing practice (GMP)
implementation of end-to-end continuous purification processes.

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{\label{works-cited}}

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\selectlanguage{english}
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\end{document}

